Co-production of biofuels and glycols

ABSTRACT

Methods and systems for co-producing higher hydrocarbons and glycols from bio-based feedstocks containing carbohydrates are disclosed.

The present application claims the benefit of U.S. Provisional PatentApplication Ser. No. 61/496,681, filed Jun. 14, 2011 the entiredisclosure of which is hereby incorporated by reference.

FIELD OF THE INVENTION

The invention relates to the production of higher hydrocarbons suitablefor use in transportation fuels and industrial chemicals from bio-basedfeedstocks.

BACKGROUND OF THE INVENTION

A significant amount of effort has been placed on developing new methodsand systems for providing energy from resources other than fossil fuels.Bio-based feedstocks are a resource that show promise as a renewablealternative source of hydrocarbons for producing fuel and chemicals.

Bio-based feedstocks including carbohydrates and “biomass” are materialsderived from living or recently living biological materials. One type ofbiomass is cellulosic biomass. Cellulosic biomass is the most abundantsource of carbohydrate in the world due to the lignocellulosic materialscomposing the cell walls. The ability to convert biomass to fuels,chemicals, energy and other materials is expected to strengthen theeconomy, minimize dependence on oil and gas resources, reduce air andwater pollution, and decrease the net rate of carbon dioxide production.

There are many challenges to overcome in developing processes ofconverting carbohydrates to higher hydrocarbons suitable for use intransportation fuels. For example, the processes used are costly andcomplex making it difficult to compete with the use of traditionalresources, such as fossil fuels. U.S. Patent Application Publication No.2007/0142633 (Yao et al.) refers to a process for the conversion ofcarbohydrates to higher hydrocarbons. An ion-exchange resin is providedto convert the carbohydrates into usable reaction products. The reactionproducts are hydrogenated, and then contacted with a zeolite containingcatalyst to form higher hydrocarbons. The conversion of carbohydrates tohydrocarbons in this system results in increased loss of desirablehydrocarbon products due to the formation of unwanted byproducts, suchas coke, carbon dioxide, and carbon monoxide. Thus, another challengefor promoting and sustaining the use of biomass is the need to eliminatethe formation of undesirable byproducts such as carbon monoxide, carbondioxide, and coke. A further challenge is to complete the conversion ofcarbohydrates to higher hydrocarbons in a limited number of steps, toobtain high yields with minimal capital investment.

Current methods for converting sugars to fuel proceed through abiological route, such as yeast fermentation, to ethanol. However,ethanol does not have a high energy density when compared to standardtransportation fuels. Currently, there is a need for the creation ofliquid biofuels of greater energy density than ethanol, which can makeuse of existing fuel infrastructure. Moreover, what is needed is amethod and system that provides efficient and high yield production ofhigher hydrocarbons suitable for use in transportation fuels andindustrial chemicals from bio-based feedstocks while avoiding orminimizing the production of unwanted by-products.

Efficient conversion of carbohydrates to glycols such as ethylene glycolor propylene glycol at high yield for use as chemical products orintermediates, has been limited by the further reaction of glycols tomonooxygenates and ultimately alkanes, such that high yields of glycolscannot be obtained at high conversions of the feed carbohydrates. Use ofhigh temperatures to increase conversion in a single reaction step canalso lead to heavy ends byproduct formation from carbohydrate feeds suchas biomass or soluble sugars, due to reactivity of sugar intermediatesto form caramelans and tars, prior to stabilization by hydrogen.Alternate use of low conversion at lower temperatures in a singlereaction step requires separation and recycle of a large stream ofunconverted carbohydrates, which increases processing costs for aprocess which targets only glycols as the principal commercial product.Low conversion to maximize glycol yields in a single reaction steps alsoresults in the presence of unconverted feed carbohydrates includingsugars and sugar alcohols, and polyoxygenated species containing morethan three oxygens. These components cause excessive coke formation anddeactivation of condensation-oligomerization catalyst, upon attemptedprocessing of the monooxygenates-rich stream obtained after glycolsseparation, to liquid fuels.

SUMMARY OF THE INVENTION

An efficient process is provided which can produce a portion of glycolsfrom carbohydrates without excessive heavy ends or light alkanesbyproduct formation, and after separation of a glycol-rich stream,continue the processing for efficient conversion of amonooxygenates-rich stream to liquid fuels, to obtain high yields ofglycols and liquid fuels, with minimal formation of heavy ends, tars,and light alkane (less than C₅) gaseous byproducts.

In an embodiment, a method comprises providing a bio-based feedstockstream containing carbohydrates and water; contacting, in a firstreaction system, the bio-based feedstock stream with hydrogen in thepresence of a hydrogenolysis catalyst at a temperature in the range of120° C. to 280° C. and 0.1 to 150 bar of hydrogen to produce ahydrogenolysis stream containing at least 5 wt %, based on the totaloxygenates content, of glycols that comprises ethylene glycol (EG) and1,2-propylene glycol (PG), and other monooxygenates; contacting, in asecond reaction system, at least a first portion of said hydrogenolysisstream with hydrogen in the presence of a hydrogenolysis catalyst at atemperature in the range of 160° C. to 280° C. and in the presence of0.1 to 150 bar hydrogen to produce a oxygenated intermediate streamcomprising at least 5 wt %, based on the total oxygenates, ofmonooxygenated hydrocarbons of chain length less than 6 carbons;processing at least a portion of the oxygenated intermediate stream toform a liquid fuel; providing a second portion of said hydrogenolysisstream to a first separation system; separating a portion of saidhydrogenoloysis stream, in the first separation system, to amonooxygenates stream comprising monooxygenates and a glycol rich streamcomprising at least 10 wt %, based on the total oxygenates, of glycolsby flashing; and recovering glycols from the glycol rich stream.

In yet another embodiment, a method comprises providing a bio-basedfeedstock; contacting the bio-based feedstock with a digestive solventto provide a bio-based feedstock stream containing carbohydrates andwater; contacting, in a first reaction system, the bio-based feedstockstream with hydrogen in the presence of a hydrogenolysis catalyst at atemperature in the range of 120° C. to 280° C. and 0.1 to 150 bar ofhydrogen to produce a hydrogenolysis stream containing at least 5 wt %,based on the total oxygenates content, of glycols that comprisesethylene glycol (EG) and 1,2-propylene glycol (PG), and othermonooxygenates; contacting, in a second reaction system, at least afirst portion of said hydrogenolysis stream with hydrogen in thepresence of a hydrogenolysis catalyst at a temperature in the range of160° C. to 280° C. and in the presence of 0.1 to 150 bar hydrogen toproduce a oxygenated intermediate stream comprising at least 5 wt %,based on the total oxygenates, of monooxygenated hydrocarbons of chainlength less than 6 carbons; processing at least a portion of theoxygenated intermediate stream to form a liquid fuel; providing a secondportion of said hydrogenolysis stream to a first separation system;separating a portion of said hydrogenoloysis stream, in the firstseparation system, to a monooxygenates stream comprising monooxygenatesand a glycol rich stream comprising at least 10 wt %, based on the totaloxygenates, of glycols by flashing; and recovering glycols from theglycol rich stream.

In yet another embodiment, a method comprises providing a bio-basedfeedstock stream containing carbohydrates and water; contacting, in afirst reaction zone, the bio-based feedstock stream with hydrogen in thepresence of a hydrogenolysis catalyst at a temperature in the range of120° C. to 280° C. and 0.1 to 150 bar of hydrogen to produce ahydrogenolysis intermediate containing at least 5 wt %, based on thetotal oxygenates content, of glycols that comprises ethylene glycol (EG)and 1,2-propylene glycol (PG), and other monooxygenates; contacting, ina second reaction zone, said hydrogenolysis intermediate with hydrogenin the presence of a hydrogenolysis catalyst at a temperature in therange of 160° C. to 280° C. and in the presence of 0.1 to 150 barhydrogen to produce a combined glycol and oxygenated intermediate streamcomprising at least 5 wt %, based on the total oxygenates, ofmonooxygenated hydrocarbons of chain length less than 6 carbons, andgreater than 5 wt % glycols; separating, by flashing, the combinedglycol and oxygenated intermediate stream into a glycol rich streamcomprising at least 10 wt %, based on the total oxygenates, of glycols,and a mono-oxygenates-rich stream; processing at least a portion of themonooxygenates-rich stream to form a liquid fuel; and recovering glycolsfrom the glycol rich stream.

BRIEF DESCRIPTION OF THE DRAWINGS

These drawings illustrate certain aspects of some of the embodiments ofthe invention, and should not be used to limit or define the invention.

FIG. 1 is a process flow diagram of an embodiment of a biofuels andglycols production process of this invention.

FIG. 2 schematically illustrates a block flow diagram of an embodimentof a biofuels and glycols production process of this invention.

FIG. 3 schematically illustrates a block flow diagram of anotherembodiment of a biofuels and glycols production process of thisinvention.

FIG. 4 shows a cross plot of conversion of sorbitol, vs. selectivity toglycols including glycerol.

DETAILED DESCRIPTION OF THE INVENTION

The invention relates to methods and systems for producing higherhydrocarbons and glycols from bio-based feedstocks, such ascarbohydrates, which include sugars, sugar alcohols, celluloses,lignocelluloses, hemicelluloses, lignocellulosic biomass, and anycombination thereof, to form higher hydrocarbons suitable for use intransportation fuels and industrial chemicals, while minimizing theformation of undesirable by-products such as coke, carbon dioxide, andcarbon monoxide. The higher hydrocarbons produced are useful in formingtransportation fuels, such as synthetic gasoline, diesel fuel, and jetfuel, as well as industrial chemicals. The glycols produced are usefulin many conventional chemicals applications as solvents and as chemicalprecursors. As used herein, the term “higher hydrocarbons” refers tohydrocarbons having an oxygen to carbon ratio less than at least onecomponent of the bio-based feedstock. As used herein the term“hydrocarbon” refers to an organic compound comprising primarilyhydrogen and carbon atoms, which is also an unsubstituted hydrocarbon.In certain embodiments, the hydrocarbons of the invention also compriseheteroatoms (e.g., oxygen or sulfur) and thus the term “hydrocarbon” mayalso include substituted hydrocarbons. The term “soluble carbohydrates”refers to oligosaccharides and monosaccharides that are soluble in thedigestive solvent and that can be used as feedstock to thehydrogenolysis reaction (e.g., pentoses and hexoses).

In an embodiment, the reactions described below are carried out in anysystem of suitable design, including systems comprising continuous-flow,batch, semi-batch, or multi-system vessels and reactors. One or morereactions may take place in an individual vessel and the process is notlimited to separate reaction vessels for each reaction. In anembodiment, the invention utilizes a fixed or fluidized catalytic bedsystem. Preferably, the invention is practiced using a continuous-flowsystem at steady-state.

The methods and systems of the invention have the advantage ofconverting bio-based feedstocks, optionally without any additionalcostly purification steps to form higher energy density product of loweroxygen/carbon ratio including higher alkanes, olefins, and aromatics.Another advantage of the present invention includes the fact that glycolcoproducts are readily produced as intermediates and easily separated byflash distillation from the primarily mono-oxygenate intermediates usedto produce liquid biofuels by condensation or oligomerization reactions.

While not intending to be limited by theory, it is believed that somecarbohydrates may thermally degrade at the conditions needed to producehigher hydrocarbons. In addition, the inclusion of some higher polyolsin a reaction to form high hydrocarbons can result in formation of tarsor coke. As an advantage of the present process, smaller reactor volumesand catalyst charges can be employed to only partially convertcarbohydrate feeds to a mixture of primarily polyols including glycolsand monooxygenates, as it is not necessary to drive the reactions tocomplete conversion of carbohydrates, in order to protectcondensation-oligomerization catalysts used in subsequent steps, fromexcessive coke and tar formation. Unreacted carbohydrates and polyolsincluding glycols are easily separated by flash distillation, andrecycled back to the aqueous phase reforming reactor. Advantages ofspecific embodiments will be described in more detail below.

An embodiment of the invention comprises (i) providing bio-basedfeedstock stream containing carbohydrates and water; (ii) contacting thebio-based feedstock stream with hydrogen in the presence of ahydrogenolysis catalyst, in the first reaction system, at a temperaturein the range of 120° C. to 280° C. and 0.1 to 150 bar of hydrogen toproduce a hydrogenolysis stream containing at least 5 wt %, preferably10 wt %, more preferably 20 wt %, based on the total oxygenates content,of glycols that comprises ethylene glycol (EG) and 1,2-propylene glycol(PG), and other monooxygenates; (iii) contacting at least a firstportion of said hydrogenolysis stream to a second reaction systemcomprising a hydrogenolysis catalyst at a temperature in the range of160° C. to 280° C. and in the presence of 0.1 to 150 bar hydrogen toproduce a oxygenated intermediate stream comprising of 5 wt %,preferably 10 wt %, more preferably 15 wt % based on the totaloxygenates, of monooxygenated hydrocarbons of chain length less than 6carbons; (iv) processing at least a portion of the oxygenatedintermediates to form a liquid fuel; (v) providing a second portion ofsaid hydrogenolysis stream to a first separation system; (vi) separatinga portion of said hydrogenolysis stream, in the first separation system,to a monooxygenates-rich stream comprising monooxygenates, polyolsincluding some glycols, alkanes, and water, and a glycol rich streamcomprising at least 10 wt %, preferably 15 wt %, more preferably 20 wt%, most preferably 25 wt %, based on the total oxygenates content, ofglycols by flashing; (vii) recovering glycols from the glycol richstream; and (vii) optionally recycling said monooxygenates-rich streamto the first reaction system. The ratio of the hydrogenolysis streamprovided to the first separation system to the second reaction system isin the range of 1.5:1 to 10:1. At least a portion of the monooxygenatesstream from the processing system and at least a portion of themonooxygenates stream from the first separation system may be recycledto the first reaction system to digest biomass to produce the aqueouscarbohydrate feed.

Use of separate processing zones for steps (ii) and (iii) allowsconditions to be optimized for digestion and hydrogenation orhydrogenolysis of the digested biomass components, independent fromoptimization of the conversion of oxygenated intermediates to themonooxygenates-rich stream comprising monooxygenates as well as somepolyols and alkanes in addition to water, before feeding to step (iv) tomake higher hydrocarbon fuels. A lower reaction temperature in step(iii) may be advantageous to minimize heavy ends byproduct formation, byconducting the hydrogenation and hydrogenolysis steps initially at a lowtemperature, whereby sugar molecules formed upon hydrolysis ofcarbohydrate feeds, are stabilized by hydrogenation to avoid formationof caramelans and tars. This has been observed to result in anintermediates stream which is rich in diols and polyols, but essentiallyfree of heavy ends and non-hydrogenated monosaccharides which otherwisewould serve as heavy ends precursors. The subsequent conversion ofmostly solubilized and hydrogenated intermediates can be doneefficiently in a second reaction zone at a higher temperature,optionally with an alternate catalyst optimized for production ofmonooxygenated hydrocarbons suitable for processing to liquid fuels,where residence time is minimized to avoid the undesired continuedreaction of monooxygenates to form light alkane or alkene byproducts.Effective choice of catalyst can minimize the formation of light alkanesof carbon number less than five, which are not useful as liquid fuels.In this manner, overall yields to desired glycols andmonooxygenates-rich streams may be improved, via conducting theconversion in two or more stages.

FIG. 2 illustrates an embodiment of a process 100 according to thepresent invention. Another embodiment is illustrated in FIG. 3. FIG. 1is a process flow diagram of one embodiment of such process. FIG. 4describes yield of glycols including glycerol vs. conversion ofsorbitol, for typical catalysts used with this invention.

In reference to FIG. 2, in one embodiment of the invention process 100,bio-based feedstock stream containing carbohydrates and water 102 frombiomass is provided to a hydrogenolysis system (first reaction system orzone) 106 containing a hydrogenolysis catalyst whereby the carbohydrateis catalytically reacted with hydrogen 104 in the presence of ahydrogenolysis catalyst at a temperature in the range of 120° C. to 280°C. and in the range of 0.1 to 150 bar of hydrogen to produce ahydrogenolysis stream containing at least 5 wt %, preferably at least 10wt % based on the total oxygenates content, of glycols, that comprisesethylene glycol (EG) and 1,2-propylene glycol (PG), and othermonooxygenates 108, and at least a first portion of the hydrogenolysisstream 110 is provided to a second reaction system (or zone) 120containing a hydrogenolysis catalyst whereby the hydrogenolysis streamis catalytically reacted with hydrogen 122 at a temperature in the rangeof 160° C. to 280° C. and in the range of 0.1 to 150 bar hydrogen toproduce an oxygenated intermediate stream 124 containing at least 5 wt%, based on the total oxygenates content, of monooxygenated hydrocarbonsof chain length less than 6 carbons. Then at least a portion of theoxygenated intermediate stream 124 is provided to a processing system150 to produce higher hydrocarbons to form a liquid fuel 152. A portionof a second portion of the hydrogenolysis stream 112 is optionallyprovided as a recycle stream to the hydrogenolysis system. A secondportion of the hydrogenolysis stream 114 is provided to a firstseparation system 206 such as a light ends column, that removes a smallportion of monooxygenates formed in the hydrogenolysis system with waterby flashing to provide a glycol enriched stream 210 andmonooxygenates-rich stream 208. The glycol enriched stream as a secondportion (bottoms) from the first separation system is provided to asecond separation system 220, a polyols recovery column, where EG and PGare separated overhead as glycol stream 224 and heavier streamcontaining heavier glycols and unconverted carbohydrates (sorbitol) arerecycled back to the feed 102 or hydrogenolysis system 106. The glycolstream may further be finished in a further third separation column 250to produce finished EG, PG or a mixture of EG and PG product(s) 252.Monooxygenates-rich stream 208 is also fed to processing system 150 toproduce higher hydrocarbons and aromatics to form a liquid fuel.

Oxygenated intermediate stream 124 and monooxygenates-rich stream 208(monooxygenates solvent streams) may optionally be used to providesolvent to digest biomass to produce bio-based feedstock stream 102. Inone embodiment, the oxygenated intermediate stream from the secondreaction system may be flash distilled to provide a stream containing atleast a portion of monooxygenate intermediates, and optionally a portionof the glycols, as feed for production of liquid fuels. Further, thebottoms from flashing containing enriched concentrations of unconvertedcarbohydrates, sugar alcohols, glycols, and some monooxygenates, may berecycle to the first reaction system.

In reference to FIG. 3, in one embodiment of the invention process 300,bio-based feedstock 302 from biomass is provided to a digestion system303 that may have one or more digester(s), whereby the biomass iscontacted with a digestive solvent. The digestive solvent is optionallyat least a portion recycled from the hydrogenolysis reaction as arecycle stream. The hydrogenolysis recycle stream can comprise a numberof components including in situ generated solvents, which may be usefulas digestive solvent at least in part or in entirety. The term “in situ”as used herein refers to a component that is produced within the overallprocess; it is not limited to a particular reactor for production or useand is therefore synonymous with an in-process generated component. Thein situ generated solvents may comprise oxygenated intermediates. Thenon-extractable solids may be removed from the reaction as an outletstream. The treated bio-based feedstock stream 305 is an intermediatestream that may comprise the treated biomass at least in part in theform of soluble carbohydrates and water. The composition of the treatedbio-based feedstock stream 305 may vary and may comprise a number ofdifferent compounds. Preferably, the contained carbohydrates will have 2to 12 carbon atoms, and even more preferably 2 to 6 carbon atoms. Thecarbohydrates may also have an oxygen to carbon ratio from 0.5:1 to1:1.2. Oligomeric carbohydrates containing more than 12 carbon atoms mayalso be present. At least a portion of the treated bio-based feedstockstream is provided to a hydrogenolysis system (first reaction system orzone) 306 containing a hydrogenolysis catalyst whereby the carbohydrateis catalytically reacted with hydrogen 304 in the presence of ahydrogenolysis catalyst at a temperature in the range of 120° C. to 280°C. and in the range of 0.1 to 150 bar of hydrogen to produce ahydrogenolysis stream containing at least 5 wt %, preferably at least 10wt % based on the total oxygenates content, of glycols, that comprisesethylene glycol (EG) and 1,2-propylene glycol (PG), and othermonooxygenates 308, and at least a first portion of the hydrogenolysisstream 310 is provided to a second reaction system (or zone) 320containing a hydrogenolysis catalyst whereby the hydrogenolysis streamis catalytically reacted with hydrogen 322 at a temperature in the rangeof 160° C. to 280° C. and in the range of 0.1 to 150 bar hydrogen toproduce a oxygenated intermediate stream 324 containing at least 5 wt %,based on the total oxygenates content, of monooxygenated hydrocarbons ofchain length less than 6 carbons. Then at least a portion of theoxygenated intermediate stream 324 is provided to a processing system350 to produce higher hydrocarbons to form a liquid fuel 352. A portionof a second portion of the hydrogenolysis stream 312 may optionally beprovided as a recycle stream to the hydrogenolysis system or digestionsystem. A second portion of the hydrogenolysis stream 314 is provided toa first separation system 406 such as a light ends column, that removesa small portion of monooxygenates formed in the hydrogenolysis systemwith water by flashing to provide a glycol enriched stream 410 andmonooxygenates stream 408. The glycol enriched stream as a secondportion (bottoms) from the first separation system is provided to asecond separation system 420, a polyols recovery column, where EG and PGare separated overhead as glycol stream 424 and heavier streamcontaining heavier glycols and unconverted carbohydrates (sorbitol) arerecycled back to the feed 302 or digestion system 303 or hydrogenolysissystem 306. The glycol stream may further be finished in a further thirdseparation column 450 to produce finished EG, PG or a mixture of EG andPG product(s) 452. Monooxygenates stream 408 may also fed to processingsystem 350 to produce higher hydrocarbons and aromatics to form a liquidfuel. Oxygenated intermediate stream 324 and monooxygenates stream 408(monooxygenates solvent streams) may optionally be used to providesolvent to digest biomass as recycle stream. A combined stream 324 and408 may be flashed to produce an overhead solvent concentrate forrecycle. The treated bio-based feedstock stream 305 may be optionallywashed prior to feeding to the hydrogenolysis system (or zone) 306. Ifwashed, water is most typically used as wash solvent. The digestionsystem 303 and the hydrogenolysis system (or zone) 306 may be carriedout in a separate or the same reactor.

In another embodiment (not shown), hydrogenolysis intermediate from afirst reaction zone is routed in its entirety to a second reaction zoneoperating at a temperature at least ten degrees Celsius higher than thefirst reaction zone to produce a combined glycol and monooxygenates-richstream. The reaction zones may comprise two zones within a singlereactor. A portion of ethylene glycol and propylene glycol are recoveredfrom the combined stream, and at least a portion of themonooxygenates-rich stream is processed into liquid fuels. Optionally,either or both of the glycols rich stream or monooxygenates stream arerecycled as solvent to the first reaction zone, or optionally to adigester to solubilize biomass or bio-based feeds for supply to thefirst reaction zone

In yet another embodiment, providing a bio-based feedstock streamcontaining carbohydrates and water, contacting, in a first reactionzone, the bio-based feedstock stream with hydrogen in the presence of ahydrogenolysis catalyst at a temperature in the range of 120° C. to 280°C. and 0.1 to 150 bar of hydrogen to produce a hydrogenolysisintermediate containing at least 5 wt %, based on the total organicscontent, of glycols comprising ethylene glycol (EG) and 1,2-propyleneglycol (PG), and other organics; routing the hydrogenolysis intermediatewith glycols to a second reaction zone operating at a temperature atleast 10° C. higher than the first reaction zone, separating at least aportion of said monooxygenates and glycol stream by flashing, to aglycol rich stream comprising at least 10 wt %, based on the totalorganics, of glycols and a monooxygenates-rich stream comprising otherorganics and water; optionally recycling a portion of either or both ofthe glycols rich stream or monooxygenates-rich stream to serve assolvent for the initial reaction step, processing at least a portion ofthe monooxygenates-rich stream to form a liquid fuel; and recoveringglycols from the glycol rich stream. The recycled solvent stream may beoptionally recycled to a digester, to provide digested biomass-feed forthe initial reaction zone.

In an embodiment, the carbohydrates fed to the process as bio-basedfeedstock stream may be derived from an organic source (e.g., sugars andstarches from corn or sugarcane). In another embodiment, the bio-basedfeedstock streams are derived from bio-based feedstocks. Bio-basedfeedstocks can include biomass. As used herein, the term “biomass” meansorganic materials produced by plants (e.g., leaves, roots, seeds andstalks), and microbial and animal metabolic wastes. Common sources ofbiomass include: agricultural wastes (e.g., corn stalks, straw, seedhulls, sugarcane leavings, bagasse, nutshells, and manure from cattle,poultry, and hogs); wood materials (e.g., wood or bark, sawdust, timberslash, and mill scrap); municipal waste, (e.g., waste paper and yardclippings); and energy crops, (e.g., poplars, willows, switch grass,alfalfa, prairie bluestream, corn, and soybean). The term “biomass” alsorefers to the primary building blocks of all the above, including, butnot limited to, saccharides, lignins, celluloses, hemicelluloses, andstarches. Any suitable (e.g., inexpensive and/or readily available) typeof biomass can be used. Suitable lignocellulosic biomass can be, forexample, selected from, but not limited to, forestry residues,agricultural residues, herbaceous material, municipal solid wastes,waste and recycled paper, pulp and paper mill residues, and combinationsthereof. Thus, in some embodiments, the biomass can comprise, forexample, corn stover, straw, bagasse, miscanthus, sorghum residue,switch grass, bamboo, water hyacinth, hardwood, hardwood chips, hardwoodpulp, softwood, softwood chips, softwood pulp, and/or combination ofthese feedstocks. The biomass can be chosen based upon a considerationsuch as, but not limited to, cellulose and/or hemicelluloses content,lignin content, growing time/season, growing location/transportationcost, growing costs, harvesting costs and the like. Plant materialsstore carbohydrates either as sugars, starches, celluloses,lignocelluloses, hemicelluloses, and any combination thereof. In oneembodiment, the carbohydrates include monosaccharides, polysaccharidesor mixtures of monosaccharides and polysaccharides. As used herein, theterm “monosaccharides” refers to hydroxy aldehydes or hydroxy ketonesthat cannot be hydrolyzed to smaller units. Examples of monosaccharidesinclude, but are not limited to, dextrose, glucose, fructose andgalactose. As used herein, the term “polysaccharides” refers tosaccharides comprising two or more monosaccharide units. Examples ofpolysaccharides include, but are not limited to, sucrose, maltose,cellobiose, cellulose and lactose.

Prior to treatment with the digestive solvent, the untreated biomass canbe washed and/or reduced in size (e.g., chopping, crushing or debarking)to a convenient size and certain quality that aids in moving the biomassor mixing and impregnating the chemicals from digestive solvent. Thus,in some embodiments, providing biomass can comprise harvesting alignocelluloses-containing plant such as, for example, a hardwood orsoftwood tree. The tree can be subjected to debarking, chopping to woodchips of desirable thickness, and washing to remove any residual soil,dirt and the like.

In the digestion system, the size-reduced biomass is contacted with thedigestive solvent in at least one digester where the digestion reactiontakes place. The digestive solvent must be effective to digest lignins.

In one aspect of the embodiment, the digestive solvent maybe aKraft-like digestive solvent that contains (i) at least 0.5 wt %,preferably at least 4 wt %, to at most 20 wt %, more preferably to 10 wt%, based on the digestive solvent, of at least one alkali selected fromthe group consisting of sodium hydroxide, sodium carbonate, sodiumsulfide, potassium hydroxide, potassium carbonate, ammonium hydroxide,and mixtures thereof, (ii) optionally, 0 to 3%, based on the digestivesolvent, of anthraquinone, sodium borate and/or polysulfides; and (iii)water (as remainder of the digestive solvent). In some embodiments, thedigestive solvent may have an active alkali of between 5 to 25%, morepreferably between 10 to 20%. The term “active alkali”(AA), as usedherein, is a percentage of alkali compounds combined, expressed assodium oxide based on weight of the biomass less water content (drysolid biomass). If sodium sulfide is present in the digestive solvent,the sulfidity can range from about 15% to about 40%, preferably fromabout 20 to about 30%. The term “sulfidity”, as used herein, is apercentage ratio of Na₂S, expressed as Na₂O, to active alkali. Digestivesolvent to biomass ratio can be within the range of 0.5 to 50,preferably 2 to 10. The digestion is carried out typically at acooking-liquor to biomass ratio in the range of 2 to 6, preferably 3 to5. The digestion reaction is carried out at a temperature within therange of from about 60° C., preferably 100° C., to about 230° C., and aresidence time within 0.25 h to 24 h. The reaction is carried out underconditions effective to provide a pretreated bio-based feedstock stream,and a chemical liquor stream containing alkali compounds and dissolvedlignin and hemicelluloses material.

The digester can be, for example, a pressure vessel of carbon steel orstainless steel or similar alloy. The digestion system can be carriedout in the same vessel or in a separate vessel. The cooking can be donein continuous or batch mode. Suitable pressure vessels include, but arenot limited to the “PANDIA™ Digester” (Voest-Alpine IndustrienlagenbauGmbH, Linz, Austria), the “DEFIBRATOR Digester” (Sunds Defibrator ABCorporation, Stockholm, Sweden), M&D (Messing & Durkee) digester (BauerBrothers Company, Springfield, Ohio, USA) and the KAMYR Digester(Andritz Inc., Glens Falls, N.Y., USA). The digestive solvent has a pHfrom 10 to 14, preferably around 12 to 13 depending on the concentrationof active alkali AA. The contents can be kept at a temperature withinthe range of from 100° C. to 230° C. for a period of time, morepreferably within the range from about 130° C. to about 180° C. Theperiod of time can be from about 0.25 to 24.0 hours, preferably fromabout 0.5 to about 2 hours, after which the pretreated contents of thedigester are discharged. For adequate penetration, a sufficient volumeof liquor is required to ensure that all the biomass surfaces arewetted. Sufficient liquor is supplied to provide the specified digestivesolvent to biomass ratio. The effect of greater dilution is to decreasethe concentration of active chemical and thereby reduce the reactionrate.

In a system using the digestive solvent such as a Kraft-like digestivesolvent similar to those used in a Kraft pulp and paper process, thechemical liquor may be regenerated in a similar manger to a Kraft pulpand paper chemical regeneration process. For example, in reference toFIG. 3 when used in a Kraft-like digestive solvent system, therecaustisized chemical recycle stream obtained by regenerating at leasta portion of the solvent liquor stream through a chemical regenerationsystem. In an embodiment, chemical liquor stream is obtained byconcentrating at least a portion of the solvent liquor stream in aconcentration system thereby producing a concentrated chemical liquorstream then burning the concentrated chemical liquor stream in a boilersystem thereby producing chemical recycle stream and a flue gas streamthen converting the sodium containing compounds to sodium hydroxide inthe recaustisizing system by contacting with lime (CaO) producing therecaustisized chemical recycle stream that can be used as a portion ofthe digestive solvent containing sodium hydroxide.

In another embodiment, an at least partially water miscible organicsolvent that has partial solubility in water, preferably greater than 2weight percent in water, may be used as digestive solvent to aid indigestion of lignin, and the nitrogen, and sulfur compounds. In one suchembodiment, the digestive solvent is a water-organic solvent mixturewith optional inorganic acid promoters such as HCl or sulfuric acid.Organic solvents exhibiting full or partial water solubility arepreferred digestive solvents. In such a process, the organic digestivesolvent mixture can be, for example, methanol, ethanol, acetone,ethylene glycol, triethylene glycol and tetrahydrofurfuryl alcohol.Organic acids such as acetic, oxalic, acetylsalicylic and salicylicacids can also be used as catalysts (as acid promoter) in the at leastpartially miscible organic solvent process. Temperatures for thedigestion may range from about 130 to about 220 degrees Celsius,preferably from about 140 to 180 degrees Celsius, and contact times from0.25 to 24 hours, preferably from about one to 4 hours. Preferably, apressure from about 1 to 80 bar, and most typically from about 5 to 50bar, is maintained on the system to avoid boiling or flashing away ofthe solvent.

Optionally the pretreated biomass stream can be washed prior tohydrogenolysis. In the wash system, the pretreated biomass stream can bewashed to remove one or more of non-cellulosic material, and non-fibrouscellulosic material prior to hydrogenolysis. The pretreated biomassstream is optionally washed with a water stream under conditions toremove at least a portion of lignin, hemicellulosic material, and saltsin the pretreated biomass stream. For example, the pretreated biomassstream can be washed with water to remove dissolved substances,including degraded, but non-processable cellulose compounds, solubilisedlignin, and/or any remaining alkaline chemicals such as sodium compoundsthat were used for cooking or produced during the cooking (orpretreatment). The washed pretreated biomass stream may contain highersolids content by further processing such as mechanical dewatering asdescribed below.

In a preferred embodiment, the pretreated biomass stream is washedcounter-currently. The wash can be at least partially carried out withinthe digester and/or externally with separate washers. In one embodimentof the invention process, the wash system contains more than one washsteps, for example, first washing, second washing, third washing, etc.that produces washed pretreated biomass stream from first washing,washed pretreated biomass stream from second washing, etc. operated in acounter current flow with the water, that is then sent to subsequentprocesses as washed pretreated biomass stream. The water is recycledthrough first recycled wash stream and second recycled wash stream andthen to third recycled wash stream. Water recovered from the chemicalliquor stream by the concentration system can be recycled as wash waterto wash system. It can be appreciated that the washed steps can beconducted with any number of steps to obtain the desired washedpretreated biomass stream. Additionally, the washing may adjust the pHfor subsequent steps where the pH is about 2.0 to 10.0, where optimal pHis determined by the catalyst employed in the hydrogenolysis step. Basessuch as alkali base may be optionally added, to adjust pH.

In the embodiment shown in FIG. 2 and FIG. 3, the bio-based feedstocksare optionally reacted in a hydrogenation reaction and then ahydrogenolysis reaction to form suitable stable hydroxyl intermediatesthat comprise glycols for further processing. In an embodiment of theinvention, the hydrogenation reaction is optional and the hydrogenolysisreaction alone is suitable to form the desired glycols. In anotherembodiment of the invention, the carbohydrates are passed through thehydrogenolysis reaction prior to being passed through the hydrogenationreaction (thus hydrogenolysis reaction and hydrogenation reaction arereversed from the order). In an embodiment of the invention, thehydrogenation and hydrogenolysis reactions occur in the same vessel togenerate glycols and other stable oxygenated intermediates to be fedinto a processing reaction and separation. In an embodiment, waterremoval could be conducted prior to the hydrogenolysis reaction.

In one embodiment of the invention, the bio-based feedstock isoptionally first hydrolyzed in a liquid medium such as an aqueoussolution or aqueous solution with organic solvent (e.g., a recycledportion of the monooxygenates), to obtain an aqueous carbohydrate streamfor use in the process. Various biomass hydrolysis methods may besuitable, including, but not limited to, acid hydrolysis, alkalinehydrolysis, enzymatic hydrolysis, and hydrolysis using hot-compressedwater. In certain embodiments, the hydrolysis reaction can occur at atemperature between 100° C. and 250° C. and pressure between about 100kPa and 10,000 kPa. In embodiments including strong acid and enzymatichydrolysis, the hydrolysis reaction can occur at temperatures as low asambient temperature and pressure between 100 kPa and 10,000 kPa. In someembodiments, the hydrolysis reaction may comprise a hydrolysis catalyst(e.g., a metal or acid catalyst) to aid in the hydrolysis reaction. Thehydrolysis catalyst can be any catalyst capable of effecting ahydrolysis reaction. For example, suitable hydrolysis catalysts include,but are not limited to, acid catalysts, base catalysts, metal catalysts,and any combination thereof. Acid catalysts can include organic acidssuch as acetic acid, formic acid, and levulinic acid. Base catalysts caninclude caustics. In an embodiment, the acid catalyst can be generatedas a byproduct during the hydrogenation and/or hydrogenolysis reactions.In certain embodiments, the hydrolysis of the bio-based feedstocks canoccur in conjunction with the hydrogenation and/or hydrogenolysisreactions. In such embodiments, a co-catalyst or catalytic support maybe added to the hydrogenation and/or hydrogenolysis reactions tofacilitate the hydrolysis reaction.

Various factors affect the conversion of the bio-based feedstock in thehydrolysis reaction. In some embodiments, hemi-cellulose can beextracted from the bio-based feedstock within an aqueous fluid andhydrolyzed at temperatures below 160° C. to produce a C5 carbohydratefraction. At increasing temperatures, this C5 fraction can be thermallydegraded. It is therefore advantageous to convert the C5, C6, or othersugar intermediates directly into more stable intermediates such assugar alcohols (monooxygenates). By recycling a portion of themonooxygenates from the hydrogenation and/or hydrogenolysis reactionsand performing additional biomass hydrolysis with this recycled liquid,the concentration of active stable hydroxyl intermediates can beincreased to commercially viable concentrations without water dilution.Typically, a concentration of at least 2%, or 5% or preferable greaterthan 8% of total organic intermediates (e.g., the recycled stablehydroxyl intermediates plus the hydrolyzed carbohydrates) in water, maybe suitable for a viable process. This may be determined by sampling theintermediate stream at the outlet of the hydrolysis reaction and using asuitable technique such as chromatography to identify the concentrationof total organics.

Cellulose extraction begins above 160° C., with solubilization andhydrolysis becoming complete at temperatures around 190° C., aided byorganic acids (e.g., carboxylic acids) formed from partial degradationof carbohydrate components. Some lignins can be solubilized beforecellulose, while other lignins may persist to higher temperatures.Organic in situ generated solvents, which may comprise a portion of themonooxygenates, including, but not limited to, light alcohols andglycols, can assist in solubilization and extraction of lignin and othercomponents.

At temperatures ranging from 250° C. to 275° C., carbohydrates candegrade through a series of complex self-condensation reactions to formcaramelans, which are considered degradation products that are difficultto convert to fuel products. In general, some degradation reactions canbe expected with aqueous reaction conditions upon application oftemperature, given that water will not completely suppressoligomerization and polymerization reactions.

The temperature of the hydrolysis reaction can be chosen so that themaximum amount of extractable carbohydrates are hydrolyzed and extractedas carbohydrates from the bio-based feedstock while limiting theformation of degradation products. In some embodiments, a plurality ofreactor vessels may be used to carry out the hydrolysis reaction. Thesevessels may have any design capable of carrying out a hydrolysisreaction. Suitable reactor vessel designs can include, but are notlimited to, co-current, counter-current, stirred tank, and/or fluidizedbed reactors. In this embodiment, the bio-based feedstock may first beintroduced into a reactor vessel operating at approximately 160° C. Atthis temperature the hemicellulose may be hydrolyzed to extract the C5carbohydrates and some lignin without degrading these products. Theremaining bio-based feedstock solids may then exit the first reactorvessel and pass to a second reactor vessel. The second vessel may beoperated between 160° C. and 250° C. so that the cellulose is furtherhydrolyzed to form C6 carbohydrates. The remaining bio-based feedstocksolids may then exit the second reactor as a waste stream while theintermediate stream from the second reactor can be cooled and combinedwith the intermediate stream from the first reactor vessel. The combinedoutlet stream may then pass to the hydrogenation and/or hydrogenolysisreactors. In another embodiment, a series of reactor vessels may be usedwith an increasing temperature profile so that a desired carbohydratefraction is extracted in each vessel. The outlet of each vessel can thenbe cooled prior to combining the streams, or the streams can beindividually fed to the hydrogenation/and or hydrogenolysis reaction forconversion of the intermediate carbohydrate streams to one or morestable hydroxyl intermediate streams.

In another embodiment, the hydrolysis reaction may take place in asingle vessel. This vessel may have any design capable of carrying out ahydrolysis reaction. Suitable reactor vessel designs can include, butare not limited to, co-current, counter-current, stirred tank, orfluidized bed reactors. In some embodiments, a counter-current reactordesign is used in which the bio-based feedstock flows counter-current tothe aqueous stream, which may comprise an in situ generated solvent. Inthis embodiment, a temperature profile may exist within the reactorvessel so that the temperature within the hydrolysis reaction media ator near the bio-based feedstock inlet is approximately 160° C. and thetemperature near the bio-based feedstock outlet is approximately 200° C.to 250° C. The temperature profile may be obtained through theintroduction of an aqueous fluid comprising an in situ generated solventabove 200° C. to 250° C. near the bio-based feedstock outlet whilesimultaneously introducing a bio-based feedstock at 160° C. or below.The specific inlet temperature of the aqueous fluid and the bio-basedfeedstock will be determined based on a heat balance between the twostreams. The resulting temperature profile may be useful for thehydrolysis and extraction of cellulose, lignin, and hemicellulosewithout the substantial production of degradation products.

In one embodiment, the conversion of carbohydrate or sugar alcohol inthe first reaction zone is limited to less than 80%, preferably lessthan 70%, more preferably less than 60%, most preferably less than 50%,to maximize selectivities to glycols. In another embodiment, theconversion of carbohydrates and sugar alcohols to monooxygenatedcompounds across the first reaction system or zone and second reactionsystem or zone is less than 90%.

Other means may be used to establish an appropriate temperature profilefor the hydrolysis reaction and extraction of cellulose andhemicellulose along with other components such as lignin withoutsubstantially producing degradation products. For example, internal heatexchange structures may be used within one or more reaction vessels tomaintain a desired temperature profile for the hydrolysis reaction.Other structures as would be known to one of ordinary skill in the artmay also be used.

In certain embodiments, the hydrolysis reaction, hydrogenation reaction,hydrogenolysis reaction, and processing reactions described in thepresent invention could be conducted in a single step.

Each reactor vessel of the invention preferably includes an inlet and anoutlet adapted to remove the product stream from the vessel or reactor.In some embodiments, the vessel in which hydrolysis reaction or someportion of the hydrolysis reaction occurs may include additional outletsto allow for the removal of portions of the reactant stream to helpmaximize the desired product formation. A backmixed reactor (e.g., astirred tank, a bubble column, and/or a jet mixed reactor) may beemployed if the viscosity and characteristics of the partially digestedbio-based feedstock and liquid reaction media is sufficient to operatein a regime where bio-based feedstock solids are suspended in an excessliquid phase (as opposed to a stacked pile digester).

Further, yields of glycols and liquid fuels and energy efficiency areimproved by staging the reaction temperatures from a lower temperature(in the range of 120 to about 230° C., most preferably in the range of180 and 215° C. in the first reaction system or zone, where carbohydratefeed is converted to hydrogenolysis intermediate, followed by a secondreaction system or zone operated in the range of 220 to 250° C. wherethe monooxygenates-rich stream is produced. Complex carbohydrate feedcomponents must first be hydrolyzed with water to form monomeric sugarscontaining carbonyl or aldehydic functionality. These reactive groupsare readily hydrogenated at a low temperature (above about 100° C.), tostabilize against self condensation to form caramelan or heavy endstars. Use of lower temperature for the first reaction system or zone,minimizes heavy ends formation relative to hydrogenation, andhydrogenolysis to form the desired glycol intermediates.

We have discovered that the further reaction of glycols and polyols toform monooxygenates occurs at a slower rate average, and is advantagedvia use of a higher temperature for the second reaction system or zone.The hydrogenation and hydrogenolysis reaction are exothermic, such thataddition of process heat from external sources is not required, toobtain an increase in the steady state average temperature from thefirst reaction system or zone, to the second reaction system or zone.Use of higher average steady state temperatures in the second reactionzone also insures complete hydrogenation of carbohydrates andderivatives containing more than two oxygen atoms per molecule, as suchthat the monooxygenates-rich stream processed to liquid fuels has ahigher effective hydrogen-to-carbon ratio, defined (G. W. Huber et al.Energy Environ. Sci., 2011, 4, 2297) as (H/C)_(eff)=(H−2O)/C. Higher(H/C)_(eff) provides reduced coking for acidic catalysts used to convertthe monooxygenates-rich stream to liquid hydrocarbon fuels. Use of toohigh a temperature in the second reaction system or zone, leads toformation of alkanes from monooxygenates. While C₅+ alkanes can beincorporated directly into liquid fuels, there is a desire not toconvert C₄− monooxygenates into the corresponding alkanes, as themonooxygenates can be dimerized or trimerized to a desirable fuelcomponent in the fuel conversion step, whereas alkanes are nonreactive,such that C₄− fraction is lost as byproduct gas.

It is therefore advantageous to operate the second reaction system orzone at a higher maximum or average steady state reaction temperaturethan the first reaction system or zone, to minimize heavy endsformation, maximize yields of glycols and liquid fuels, with minimaldeactivation of catalysts due to coke deposition, and minimization ofthe amount of process energy required to be added from external sources.

It is understood that in one embodiment, the biomass does not need to behydrolyzed, as the carbohydrate containing biomass may already be insuitable aqueous form (e.g., raw cane juice concentrate) for convertingthe bio-based feedstock to higher hydrocarbons.

In an embodiment of the invention, the intermediate carbohydrate streamproduced by the hydrolysis reaction may be converted to stable hydroxylintermediates including, but not limited to, glycols, and alcohols. Ingeneral, without being limited by any particular theory, a suitableconversion reaction or reactions can include, without limitation:hydrogenolysis, consecutive hydrogenation-hydrogenolysis, consecutivehydrogenolysis-hydrogenation, and combined hydrogenation-hydrogenolysisreactions, resulting in the formation of monooxygenates that can beeasily converted to higher hydrocarbons by one or more processingreactions.

In an embodiment of the invention, it is desirable to convert thecarbohydrates and optionally a portion of monooxygenates to smallermolecules that will be more readily converted to desired higherhydrocarbons. A suitable method for this conversion is through ahydrogenolysis reaction.

Various processes are known for performing hydrogenolysis ofcarbohydrates. One suitable method includes contacting a carbohydrate ormonooxygenates with hydrogen or hydrogen mixed with a suitable gas and ahydrogenolysis catalyst in a hydrogenolysis reaction under conditionssufficient to form a reaction product comprising smaller molecules orglycols. Most typically, hydrogen is dissolved in the liquid mixture ofcarbohydrate, which is in contact with the catalyst under conditions toprovide catalytic reaction. At least a portion of the carbohydrate feedis contacted directly with hydrogen in the presence of thehydrogenolysis catalyst. By the term “directly”, the reaction is carriedout on at least a portion of the carbohydrate without necessary stepwisefirst converting all of the carbohydrates into a stable hydroxylintermediate. As used herein, the term “smaller molecules or glycols”includes any molecule that has a lower molecular weight, which caninclude a smaller number of carbon atoms or oxygen atoms than thestarting carbohydrate. In an embodiment, the reaction products includesmaller molecules that include glycols and alcohols.

In an embodiment, a carbohydrate (e.g., a 5 and/or 6 carbon carbohydratemolecule) can be converted to monooxygenates as hydrogenolysis streamcomprising propylene glycol, ethylene glycol, and glycerol using ahydrogenolysis reaction in the presence of a hydrogenolysis catalyst.The hydrogenolysis catalyst may include Cr, Mo, W, Re, Mn, Cu, Cd, Fe,Co, Ni, Pt, Pd, Rh, Ru, Ir, Os, and alloys or any combination thereof,either alone or with promoters such as Au, Ag, Cr, Zn, Mn, Sn, Bi, B, O,and alloys or any combination thereof. The promoters may allow forhydrogenation and hydrogenolysis reactions to occur at the same time.The hydrogenolysis catalyst can also include a carbonaceous pyropolymercatalyst containing transition metals (e.g., chromium, molybdemum,tungsten, rhenium, manganese, copper, cadmium) or Group VIII metals(e.g., iron, cobalt, nickel, platinum, palladium, rhodium, ruthenium,iridium, and osmium). In certain embodiments, the hydrogenolysiscatalyst can include any of the above metals combined with an alkalineearth metal oxide or adhered to a catalytically active support. Incertain embodiments, the catalyst described in the hydrogenolysisreaction can include a catalyst support as described herein for thehydrogenation reaction. Further, a supported hydrogenolysis catalystcontaining (a) sulfur and (b) molybdenum and/or tungsten and (c) cobaltand/or nickel may be used, particularly if a sulfur poison tolerantcatalyst is preferred.

The conditions for which to carry out the hydrogenolysis reaction willvary based on the type of starting material and the desired products.One of ordinary skill in the art, with the benefit of this disclosure,will recognize the appropriate conditions to use to carry out thereaction. In general, the hydrogenolysis reaction is conducted attemperatures of at least 110° C., and preferably from at 150° C. to 240°C., and most preferably at 180° C. to 220° C. In an embodiment, thehydrogenolysis reaction is conducted under basic conditions, preferablyat a pH of 8 to 13, and even more preferably at a pH of 10 to 12. In anembodiment, the hydrogenolysis reaction is conducted at 0.1 to 150 barof hydrogen and pressures in a range between 60 kPa and 16500 kPa, andpreferably in a range between 1700 kPa and 14000 kPa, and even morepreferably between 4800 kPa and 11000 kPa.

The hydrogen used in the hydrogenolysis reaction of the currentinvention can include external hydrogen, recycled hydrogen, in situgenerated hydrogen, and any combination thereof.

In an embodiment, the use of a hydrogenolysis reaction may produce lesscarbon dioxide and a greater amount of glycols than a reaction thatresults in reforming of the reactants. For example, reforming can beillustrated by formation of isopropanol (i.e., IPA, or 2-propanol) fromsorbitol:C₆H₁₄O₆+H₂O→4H₂+3CO₂+C₃H₈O; dHR=−40 J/g-mol  (Eq. 1)

Alternately, in the presence of hydrogen, glycols and mono-oxygenatessuch as IPA can be formed by hydrogenolysis, where hydrogen is consumedrather than produced:C₆H₁₄O₆+3H₂→2H₂O+2C₃H₈O₂ ; dHR=+81 J/gmol  (Eq. 2)C₆H₁₄O₆+5H₂→4H₂O+2C₃H₈O; dHR=−339 J/gmol  (Eq. 3)

As a result of the differences in the reaction conditions (e.g.,temperatures below 250° C.), the products of the hydrogenolysis reactionmay comprise greater than 5%%, preferably 10% by mole, or alternatively,greater than 20% by mole, of glycols, which may result in a greaterconversion in a processing reaction. In addition, the use of ahydrolysis reaction rather than a reaction running at reformingconditions may result in less than 10% by mole, or alternatively lessthan 20% by mole carbon dioxide production.

In an embodiment, hydrogenolysis is conducted under neutral or acidicconditions, as needed to accelerate hydrolysis reactions in addition tothe hydrogenolysis.

In an embodiment of the invention, a hydrolyzed, substantiallyhydrolyzed, or non-hydrolyzed biomass-derived carbohydrate may convertedinto glycols and other monooxygenates comprising the correspondingalcohol derivative through a hydrogenolysis reaction in addition to anoptional hydrogenation reaction in a suitable hydrogenation reactionvessel.

The carbohydrates, glycols and other monooxygenates, from the hydrolysisreaction, or both may take place in a hydrogenation reaction to saturateone or more unsaturated bonds. Various processes are suitable forhydrogenating carbohydrates, glycols and other monooxygenates or both.One method includes contacting a feed stream with hydrogen or hydrogenmixed with a suitable gas and a catalyst under conditions sufficient tocause a hydrogenation reaction to form a hydrogenated product. Thehydrogenation catalyst generally can include a Group VIII metal and/or aGroup VI metal. Examples of such a catalyst can include, but is notlimited to, Cu, Re, Ni, Fe, Co, Ru, Pd, Rh, Pt, Os, Ir, and alloys orany combination thereof, either alone or with promoters such as W, Mo,Au, Ag, Cr, Zn, Mn, Sn, B, P, Bi, and alloys or any combination thereof.Other effective hydrogenation catalyst materials include eithersupported nickel or ruthenium modified with rhenium. In an embodiment,the hydrogenation catalyst also includes any one of the supportsdescribed below, depending on the desired functionality of the catalyst.The hydrogenation catalysts may be prepared by methods known to those ofordinary skill in the art.

In an embodiment, the hydrogenation catalyst includes a supported GroupVIII metal catalyst and a metal sponge material (e.g., a sponge nickelcatalyst). Raney nickel provides an example of an activated spongenickel catalyst suitable for use in this invention. In an embodiment,the hydrogenation reaction in the invention is performed using acatalyst comprising a nickel-rhenium catalyst or a tungsten-modifiednickel catalyst. One example of a suitable catalyst for thehydrogenation reaction of the invention is a carbon-supportednickel-rhenium catalyst.

In an embodiment, a suitable Raney nickel catalyst may be prepared bytreating an alloy of approximately equal amounts by weight of nickel andaluminum with an aqueous alkali solution, e.g., containing about 25weight % of sodium hydroxide. The aluminum is selectively dissolved bythe aqueous alkali solution resulting in a sponge shaped materialcomprising mostly nickel with minor amounts of aluminum. The initialalloy includes promoter metals (e.g., molybdenum or chromium) in theamount such that 1 to 2 weight % remains in the formed sponge nickelcatalyst. In another embodiment, the hydrogenation catalyst is preparedusing a solution of ruthenium(III) nitrosylnitrate, ruthenium (III)chloride in water to impregnate a suitable support material. Thesolution is then dried to form a solid having a water content of lessthan 1% by weight. The solid is then reduced at atmospheric pressure ina hydrogen stream at 300° C. (uncalcined) or 400° C. (calcined) in arotary ball furnace for 4 hours. After cooling and rendering thecatalyst inert with nitrogen, 5% by volume of oxygen in nitrogen ispassed over the catalyst for 2 hours.

In certain embodiments, the catalyst described includes a catalystsupport. The catalyst support stabilizes and supports the catalyst. Thetype of catalyst support used depends on the chosen catalyst and thereaction conditions. Suitable supports for the invention include, butare not limited to, carbon, silica, silica-alumina, zirconia, titania,ceria, vanadia, nitride, boron nitride, heteropolyacids, hydroxyapatite,zinc oxide, chromia, zeolites, carbon nanotubes, carbon fullerene andany combination thereof.

The catalysts used in this invention can be prepared using conventionalmethods known to those in the art. Suitable methods may include, but arenot limited to, incipient wetting, evaporative impregnation, chemicalvapor deposition, wash-coating, magnetron sputtering techniques, and thelike.

The conditions for which to carry out the hydrogenation reaction willvary based on the type of starting material and the desired products.One of ordinary skill in the art, with the benefit of this disclosure,will recognize the appropriate reaction conditions. In general, thehydrogenation reaction is conducted at temperatures of 40° C. to 250°C., and preferably at 90° C. to 200° C., and most preferably at 100° C.to 150° C. In an embodiment, the hydrogenation reaction is conducted atpressures from 500 kPa to 14,000 kPa.

In some embodiments, a plurality of reactor vessels may be used to carryout the hydrogenation reaction. The plurality of vessels may be capableof carrying out a hydrogenation reaction without producing unwantedbyproducts while minimizing degradation of wanted products. In anembodiment, the hydrogenation reaction may occur in two or more stages.In this embodiment, the bio-based feedstock may first be introduced intoa first stage reaction operating at a temperature between 40° C. to 90°C. The products may then be exposed to a second stage reaction operatingat a temperature between 80° C. to 120° C. The remaining products maythen be exposed to a third stage operating at a temperature between 120°C. and 190° C. In an embodiment, the hydrogen used in the hydrogenationreaction of the current invention can include external hydrogen,recycled hydrogen, in situ generated hydrogen, and any combinationthereof. As used herein, the term “external hydrogen” refers to hydrogenthat does not originate from the bio-based feedstock reaction itself,but rather is added to the system from another source.

In an embodiment, the invention comprises a system having a firstreactor for receiving a carbohydrate and producing a hydrogenatedproduct. Each reactor of the invention preferably includes an inlet andan outlet adapted to remove the product stream from the reactor. In anembodiment, the reactors include additional outlets to allow for theremoval of portions of the reactant stream to help maximize the desiredproduct formation, and allow for collection and recycling of by-productsfor use in other portions of the system.

In an embodiment, the invention comprises a system having a secondreactor system for receiving the hydrogenated product and converting itinto an alcohol and a polyol through a hydrolysis reaction. In certainembodiments, the hydrogenation and hydrogenolysis catalysts are the sameand may exist in the same bed in the same vessel. Each reactor of theinvention preferably includes an inlet and an outlet adapted to removethe product stream from the reactor. In an embodiment, the reactor mayinclude additional outlets to allow for the removal of portions of thereactant stream to help maximize the desired product formation, andallow for collection and recycling of by-products for use in otherportions of the system.

In an embodiment, the system of the invention includes elements thatallow for the separation of the hydrogenolysis stream into differentcomponents to promote the desired products being fed into the secondreaction system and the first separation system. For example, a suitableseparator unit includes, but is not limited to, a phase separator,stripping column, extractor, or distillation column.

In some embodiments, an outlet stream comprising at least a portion ofthe hydrogenolysis stream can pass to a processing reaction that maycomprise a condensation reaction. In an embodiment, the oxygen to carbonratio of the higher hydrocarbons produced through the condensationreaction is less than 0.5, alternatively less than 0.4, or preferablyless than 0.3 Formation of ethylene glycol and 1,2-propylene glycoloccurs during the aqueous phase reforming reaction of sugars andcarbohydrates. Continued reaction of these intermediates results in theformation of ethanol, 1-propanol and 2-propanol, which can be condensedor dehydrated and oliogomerized to higher molecular weight componentssuitable for liquid biofuels, in a subsequent reaction step. Relativerates of reaction for the various conversion steps results in anobservable concentration of ethylene glycol and propylene glycol in theintermediate reaction mixture, under appropriate conditions.

Mono-oxygenate reaction intermediates are preferred for subsequentcondensation and oligomerization steps, and are more volatile thanethylene glycol and propylene glycol, which themselves are more volatilethan sugars such as glucose, xylose, mannose, or complex sugars such assucrose, and than the corresponding sugar alcohols such as sorbitol.Mono-oxygenates and at least a fraction of water present in the reactionmixture may therefore be separated as an overheaded product via flash ormultistage distillation. Continued distillation will separate remainingwater with ethylene and propylene glycol, as a second product.Unconverted sugars and sugar alcohols, and byproduct glycerol will beconcentrated in the bottoms product from the distillation.

A single distillation column may be employed, with separation ofethylene glycol, propylene glycol and some water as a side draw stream,with removal of monooxygenates and some water as a tops stream, andsugar and sugar alcohol with some glycerol as a bottoms stream.Alternately, two distillation columns or flashers may be employed inseries, with production of monoxygenates and water as the overheadproduct from the first flasher or distillation column, and production ofethylene glycol and propylene glycol as the overhead product of a seconddistillation column, with sugars, unconverted sugar alcohols, and otherheavy ends as a bottoms product. The ethylene glycol and propyleneglycol streams may optionally be sent to a distillation column forseparation into individual components.

Light gases such as CO₂, H₂, and alkanes may be removed first byflashing, or be removed as an overhead vapor stream from the firstflasher or distillation column Flashers may entail an open vessel or apacked or trayed vessel. Distillation columns may be packed, or trayed,or employ a catalyst for further conversion of intermediates present.

A portion of the hydrogenolysis stream provided to a second reactionsystem containing a hydrogenolysis catalyst at a temperature in therange of 160 to 280° C., preferably 210 to 260° C. and in the presenceof 0.1 to 150 bar hydrogen to produce a oxygenated intermediate streamcomprising of 5 wt % of monooxygenated hydrocarbons of chain length lessthan 6 carbons, based on the total. In an embodiment, the secondreaction system is conducted under basic conditions, preferably at a pHof 8 to 13, and even more preferably at a pH of 10 to 12. In anembodiment, the hydrogenolysis reaction is conducted at pressures in arange between 60 kPa and 16500 kPa, and preferably in a range between1700 kPa and 14000 kPa, and even more preferably between 4800 kPa and11000 kPa.

It is generally preferred to operate the second reaction system or zoneat a higher temperature than the first, to facilitate conversion ofpolyoxygenates to monooxygenate components. Temperature must however belimited to prevent full hydrogenation or hydrodeoxygenation ofmonooxygenates containing less than four carbons to alkanes, tominmimize the amount of light gas byproduct. Once fully hydrogenated toalkanes, it is not possible to increase chain length in subsequentcondensation-oligomerization steps, and hence any alkane less than C₅becomes a yield loss to light gas byproduct.

Optionally, a different catalyst may be used in the second reaction zonerelative to the first reaction zone. Purification such as ion exchangeor adsorption may be used prior to the first reaction zone or the secondreaction zone, or both, to protect catalyst from impurities in the feed,or released during the hydrolysis of bio-based feeds during processing.Removal of protein derivatives such as amino acids containing N andsometimes S functional groups, may be required to prevent excessiverates of catalyst poisoning. Alternately, a poison-tolerant catalyst mebe deployed in either of the reaction zones.

The hydrogen used in the hydrogenolysis reaction of the currentinvention can include external hydrogen, recycled hydrogen, in situgenerated hydrogen, and any combination thereof.

The conditions for which to carry out the hydrogenolysis reaction willvary based on the type of starting material and the desired productsafter further process to form the liquid fuel.

At least a portion of said hydrogenolysis stream is provided to thefirst separation system such as a light ends column, that removes asmall portion of monooxygenates formed in the hydrogenolysis system withwater by flashing to provide a glycol enriched stream 210 andmonooxygenates-rich stream 208 also containing other organics. Theglycol enriched stream as a second portion (bottoms) from the firstseparation system is provided to a second separation system 220, apolyols recovery column, where EG and PG are separated overhead asglycol stream 224 and heavier stream containing heavier glyols andunconverted carbohydrates (sorbitol) are optionally recycled back to thefeed 102 or hydrogenolysis system 106. The monooxygenates-rich streamgenerally contains monooxygenates, as well as some polyoxygenatedcompounds, alkanes, and water, and the glycol rich contains at least 10wt %, preferably 15 wt %, more preferably 20 wt %, most preferably 25 wt%, based on the total oxygenates content, of glycols. The glycols can berecovered from the glycol rich stream by conventional recovery methodssuch as, for example, distillation or can be used as solvents orfeedstocks for other processes. The flashing may be carried out at atemperature in the range of about 150 to 280° C. and pressure in therange of 0.5 to 5 bar absolute.

A portion of the oxygenated intermediates can be processed to produce afuel blend in one or more processing reactions. In an embodiment, acondensation reaction can be used along with other reactions to generatea fuel blend and may be catalyzed by a catalyst comprising acid or basicfunctional sites, or both. In general, without being limited to anyparticular theory, it is believed that the basic condensation reactionsgenerally consist of a series of steps involving: (1) an optionaldehydrogenation reaction; (2) an optional dehydration reaction that maybe acid catalyzed; (3) an aldol condensation reaction; (4) an optionalketonization reaction; (5) an optional furanic ring opening reaction;(6) hydrogenation of the resulting condensation products to form a C4+hydrocarbon; and (7) any combination thereof. Acid catalyzedcondensations may similarly entail optional hydrogenation ordehydrogenation reactions, dehydration, and oligomerization reactions.Additional polishing reactions may also be used to conform the productto a specific fuel standard, including reactions conducted in thepresence of hydrogen and a hydrogenation catalyst to remove functionalgroups from final fuel product. A catalyst comprising a basic functionalsite, both an acid and a basic functional site, and optionallycomprising a metal function, may be used to effect the condensationreaction.

“Acidic” conditions or “acidic functionality” for the catalysts refer toeither Bronsted or Lewis acid acidity. For Bronsted acidity, thecatalyst is capable of donating protons (designed as H⁺) to perform thecatalytic reaction, under the conditions present in the catalyticreactor. Acidic ion exchange resins, phosphoric acid present as a liquidphase on a support, are two examples. Metal oxides such as silica,silica-aluminas, promoted zirconia or titania can provide protons H⁺associated with Bronsted acidiy in the presence of water or water vapor.Lewis acidity entails ability to accept an electron pair, and mosttypically is obtained via the presence of metal cations in a mixedmetal-oxide framework such as silica-alumina or zeolite. Determinationof acidic properties can be done via adsorption of a base such asammonia, use of indictors, or via use of a probe reaction such asdehydration of an alcohol to an olefin, which is acid catalyzed. “Basic”conditions or “base functionality” for the catalysts can refer to eitherBronsted or Lewis basicity. For Bronsted basicity, hydroxide anion issupplied by the catalyst, which may be present as an ion exchange resin,or supported liquid phase catalyst, mixed metal oxide with promoter suchas alkali, calcium, or magnesium, or in free solution. Lewis basecatalysis refers to the conditions where Lewis base catalysis is theprocess by which an electron pair donor increases the rate of a givenchemical reaction by interacting with an acceptor atom in one of thereagents or substrate (see Scott E. Denmark and Gregory L. Beutner,Lewis Base Catalysis in Organic Synthesis, Angew. Chem. Int. Ed. 2008,47, 1560-1638). Presence and characterization of basic sites for aheterogeneous catalyst may be determined via sorption of an acidiccomponent, use of probe reactions, or use of indicators. (see K. Tanabe,M. Misono, Y. Ono, H. Hattori (Eds.), New Solid Acids and Bases,Kodansha/Elsevier, Tokyo/Amsterdam, 1989, pp. 260-267). Catalysts suchas mixed metal oxides may be “amphoteric”, or capable of acting asacidic or basic catalysts depending on process conditions (pH, waterconcentration), or exhibit both acidic and basic properties underspecific operating conditions, as a result of surface structuresgenerated during formulation, or in situ during use to effect catalyticreactions

In an embodiment, the aldol condensation reaction may be used to producea fuel blend meeting the requirements for a diesel fuel or jet fuel.Traditional diesel fuels are petroleum distillates rich in paraffinichydrocarbons. They have boiling ranges as broad as 187° C. to 417° C.,which are suitable for combustion in a compression ignition engine, suchas a diesel engine vehicle. The American Society of Testing andMaterials (ASTM) establishes the grade of diesel according to theboiling range, along with allowable ranges of other fuel properties,such as cetane number, cloud point, flash point, viscosity, anilinepoint, sulfur content, water content, ash content, copper stripcorrosion, and carbon residue. Thus, any fuel blend meeting ASTM D975can be defined as diesel fuel.

The present invention also provides methods to produce jet fuel. Jetfuel is clear to straw colored. The most common fuel is anunleaded/paraffin oil-based fuel classified as Aeroplane A-1, which isproduced to an internationally standardized set of specifications. Jetfuel is a mixture of a large number of different hydrocarbons, possiblyas many as a thousand or more. The range of their sizes (molecularweights or carbon numbers) is restricted by the requirements for theproduct, for example, freezing point or smoke point. Kerosene-typeAirplane fuel (including Jet A and Jet A-1) has a carbon numberdistribution between about C8 and C16. Wide-cut or naphtha-type Airplanefuel (including Jet B) typically has a carbon number distributionbetween about C5 and C15. A fuel blend meeting ASTM D1655 can be definedas jet fuel.

In certain embodiments, both Airplanes (Jet A and Jet B) contain anumber of additives. Useful additives include, but are not limited to,antioxidants, antistatic agents, corrosion inhibitors, and fuel systemicing inhibitor (FSII) agents. Antioxidants prevent gumming and usually,are based on alkylated phenols, for example, AO-30, AO-31, or AO-37.Antistatic agents dissipate static electricity and prevent sparking.Stadis 450 with dinonylnaphthylsulfonic acid (DINNSA) as the activeingredient, is an example. Corrosion inhibitors, e.g., DCI-4A are usedfor civilian and military fuels and DCI-6A is used for military fuels.FSII agents, include, e.g., Di-EGME.

In an embodiment, the oxygenated intermediates may comprise acarbonyl-containing compound that can take part in a base catalyzedcondensation reaction. In some embodiments, an optional dehydrogenationreaction may be used to increase the amount of carbonyl-containingcompounds in the oxygenated intermediate stream to be used as a feed tothe condensation reaction. In these embodiments, the oxygenatedintermediates and/or a portion of the bio-based feedstock stream can bedehydrogenated in the presence of a catalyst.

In an embodiment, a dehydrogenation catalyst may be preferred for anoxygenated intermediate stream comprising alcohols, diols, and triols.In general, alcohols cannot participate in aldol condensation directly.The hydroxyl group or groups present can be converted into carbonyls(e.g., aldehydes, ketones, etc.) in order to participate in an aldolcondensation reaction. A dehydrogenation catalyst may be included toeffect dehydrogenation of any alcohols, diols, or polyols present toform ketones and aldehydes. The dehydration catalyst is typically formedfrom the same metals as used for hydrogenation or aqueous phasereforming, which catalysts are described in more detail above.Dehydrogenation yields are enhanced by the removal or consumption ofhydrogen as it forms during the reaction. The dehydrogenation step maybe carried out as a separate reaction step before an aldol condensationreaction, or the dehydrogenation reaction may be carried out in concertwith the aldol condensation reaction. For concerted dehydrogenation andaldol condensation, the dehydrogenation and aldol condensation functionscan be on the same catalyst. For example, a metalhydrogenation/dehydrogenation functionality may be present on catalystcomprising a basic functionality.

The dehydrogenation reaction may result in the production of acarbonyl-containing compound. Suitable carbonyl-containing compoundsinclude, but are not limited to, any compound comprising a carbonylfunctional group that can form carbanion species or can react in acondensation reaction with a carbanion species, where “carbonyl” isdefined as a carbon atom doubly-bonded to oxygen. In an embodiment, acarbonyl-containing compound can include, but is not limited to,ketones, aldehydes, furfurals, hydroxy carboxylic acids, and, carboxylicacids. The ketones may include, without limitation, hydroxyketones,cyclic ketones, diketones, acetone, propanone, 2-oxopropanal, butanone,butane-2,3-dione, 3-hydroxybutane-2-one, pentanone, cyclopentanone,pentane-2,3-dione, pentane-2,4-dione, hexanone, cyclohexanone,2-methyl-cyclopentanone, heptanone, octanone, nonanone, decanone,undecanone, dodecanone, methylglyoxal, butanedione, pentanedione,diketohexane, dihydroxyacetone, and isomers thereof. The aldehydes mayinclude, without limitation, hydroxyaldehydes, acetaldehyde,glyceraldehyde, propionaldehyde, butyraldehyde, pentanal, hexanal,heptanal, octanal, nonal, decanal, undecanal, dodecanal, and isomersthereof. The carboxylic acids may include, without limitation, formicacid, acetic acid, propionic acid, butanoic acid, pentanoic acid,hexanoic acid, heptanoic acid, isomers and derivatives thereof,including hydroxylated derivatives, such as 2-hydroxybutanoic acid andlactic acid. Furfurals include, without limitation,hydroxylmethylfurfural, 5-hydroxymethyl-2(5H)-furanone,dihydro-5-(hydroxymethyl)-2(3H)-furanone, tetrahydro-2-furoic acid,dihydro-5-(hydroxymethyl)-2(3H)-furanone, tetrahydrofurfuryl alcohol,1-(2-furyl)ethanol, hydroxymethyltetrahydrofurfural, and isomersthereof. In an embodiment, the dehydrogenation reaction results in theproduction of a carbonyl-containing compound that is combined with theoxygenated intermediates to become a part of the oxygenatedintermediates fed to the condensation reaction.

In an embodiment, an acid catalyst may be used to optionally dehydrateat least a portion of the oxygenated intermediate stream. Suitable acidcatalysts for use in the dehydration reaction include, but are notlimited to, mineral acids (e.g., HCl, H₂SO₄), solid acids (e.g.,zeolites, ion-exchange resins) and acid salts (e.g., LaCl₃). Additionalacid catalysts may include, without limitation, zeolites, carbides,nitrides, zirconia, alumina, silica, aluminosilicates, phosphates,titanium oxides, zinc oxides, vanadium oxides, lanthanum oxides, yttriumoxides, scandium oxides, magnesium oxides, cerium oxides, barium oxides,calcium oxides, hydroxides, heteropolyacids, inorganic acids, acidmodified resins, base modified resins, and any combination thereof. Insome embodiments, the dehydration catalyst can also include a modifier.Suitable modifiers include La, Y, Sc, P, B, Bi, Li, Na, K, Rb, Cs, Mg,Ca, Sr, Ba, and any combination thereof. The modifiers may be useful,inter alia, to carry out a concerted hydrogenation/dehydrogenationreaction with the dehydration reaction. In some embodiments, thedehydration catalyst can also include a metal. Suitable metals includeCu, Ag, Au, Pt, Ni, Fe, Co, Ru, Zn, Cd, Ga, In, Rh, Pd, Ir, Re, Mn, Cr,Mo, W, Sn, Os, alloys, and any combination thereof. The dehydrationcatalyst may be self supporting, supported on an inert support or resin,or it may be dissolved in solution.

In some embodiments, the dehydration reaction occurs in the vapor phase.In other embodiments, the dehydration reaction occurs in the liquidphase. For liquid phase dehydration reactions, an aqueous solution maybe used to carry out the reaction. In an embodiment, other solvents inaddition to water, are used to form the aqueous solution. For example,water soluble organic solvents may be present. Suitable solvents caninclude, but are not limited to, hydroxymethylfurfural (HMF),dimethylsulfoxide (DMSO), 1-methyl-n-pyrollidone (NMP), and anycombination thereof. Other suitable aprotic solvents may also be usedalone or in combination with any of these solvents.

In an embodiment, the processing reactions may comprise an optionalketonization reaction. A ketonization reaction may increase the numberof ketone functional groups within at least a portion of the oxygenatedintermediate stream. For example, an alcohol or other hydroxylfunctional group can be converted into a ketone in a ketonizationreaction. Ketonization may be carried out in the presence of a basecatalyst. Any of the base catalysts described above as the basiccomponent of the aldol condensation reaction can be used to effect aketonization reaction. Suitable reaction conditions are known to one ofordinary skill in the art and generally correspond to the reactionconditions listed above with respect to the aldol condensation reaction.The ketonization reaction may be carried out as a separate reactionstep, or it may be carried out in concert with the aldol condensationreaction. The inclusion of a basic functional site on the aldolcondensation catalyst may result in concerted ketonization and aldolcondensation reactions.

In an embodiment, the processing reactions may comprise an optionalfuranic ring opening reaction. A furanic ring opening reaction mayresult in the conversion of at least a portion of any oxygenatedintermediates comprising a furanic ring into compounds that are morereactive in an aldol condensation reaction. A furanic ring openingreaction may be carried out in the presence of an acidic catalyst. Anyof the acid catalysts described above as the acid component of the aldolcondensation reaction can be used to effect a furanic ring openingreaction. Suitable reaction conditions are known to one of ordinaryskill in the art and generally correspond to the reaction conditionslisted above with respect to the aldol condensation reaction. Thefuranic ring opening reaction may be carried out as a separate reactionstep, or it may be carried out in concert with the aldol condensationreaction. The inclusion of an acid functional site on the aldolcondensation catalyst may result in a concerted furanic ring openingreaction and aldol condensation reactions. Such an embodiment may beadvantageous as any furanic rings can be opened in the presence of anacid functionality and reacted in an aldol condensation reaction using abase functionality. Such a concerted reaction scheme may allow for theproduction of a greater amount of higher hydrocarbons to be formed for agiven oxygenated intermediate feed.

In an embodiment, production of a C4+ compound occurs by condensation,which may include aldol-condensation, of the oxygenated intermediates inthe presence of a condensation catalyst. Aldol-condensation generallyinvolves the carbon-carbon coupling between two compounds, at least oneof which may contain a carbonyl group, to form a larger organicmolecule. For example, acetone may react with hydroxymethylfurfural toform a C9 species, which may subsequently react with anotherhydroxymethylfurfural molecule to form a C15 species. The reaction isusually carried out in the presence of a condensation catalyst. Thecondensation reaction may be carried out in the vapor or liquid phase.In an embodiment, the reaction may take place at a temperature in therange of from about 7° C. to about 377° C., depending on the reactivityof the carbonyl group.

The condensation catalyst will generally be a catalyst capable offorming longer chain compounds by linking two molecules through a newcarbon-carbon bond, such as a basic catalyst, a multi-functionalcatalyst having both acid and base functionality, or either type ofcatalyst also comprising an optional metal functionality. In anembodiment, the multi-functional catalyst will be a catalyst having botha strong acid and a strong base functionality. In an embodiment, aldolcatalysts can comprise Li, Na, K, Cs, B, Rb, Mg, Ca, Sr, Si, Ba, Al, Zn,Ce, La, Y, Sc, Y, Zr, Ti, hydrotalcite, zinc-aluminate, phosphate,base-treated aluminosilicate zeolite, a basic resin, basic nitride,alloys or any combination thereof. In an embodiment, the base catalystcan also comprise an oxide of Ti, Zr, V, Nb, Ta, Mo, Cr, W, Mn, Re, Al,Ga, In, Co, Ni, Si, Cu, Zn, Sn, Cd, Mg, P, Fe, or any combinationthereof. In an embodiment, the condensation catalyst comprises amixed-oxide base catalysts. Suitable mixed-oxide base catalysts cancomprise a combination of magnesium, zirconium, and oxygen, which maycomprise, without limitation: Si—Mg—O, Mg—Ti—O, Y—Mg—O, Y—Zr—O, Ti—Zr—O,Ce—Zr—O, Ce—Mg—O, Ca—Zr—O, La—Zr—O, B—Zr—O, La—Ti—O, B—Ti—O, and anycombinations thereof. Different atomic ratios of Mg/Zr or thecombinations of various other elements constituting the mixed oxidecatalyst may be used ranging from about 0.01 to about 50. In anembodiment, the condensation catalyst further includes a metal or alloyscomprising metals, such as Cu, Ag, Au, Pt, Ni, Fe, Co, Ru, Zn, Cd, Ga,In, Rh, Pd, Ir, Re, Mn, Cr, Mo, W, Sn, Bi, Pb, Os, alloys andcombinations thereof. Such metals may be preferred when adehydrogenation reaction is to be carried out in concert with the aldolcondensation reaction. In an embodiment, preferred Group IA materialsinclude Li, Na, K, Cs and Rb. In an embodiment, preferred Group IIAmaterials include Mg, Ca, Sr and Ba. In an embodiment, Group IIBmaterials include Zn and Cd. In an embodiment, Group IIIB materialsinclude Y and La. Basic resins include resins that exhibit basicfunctionality. The base catalyst may be self-supporting or adhered toany one of the supports further described below, including supportscontaining carbon, silica, alumina, zirconia, titania, vanadia, ceria,nitride, boron nitride, heteropolyacids, alloys and mixtures thereof.

In one embodiment, the condensation catalyst is derived from thecombination of MgO and Al₂O₃ to form a hydrotalcite material. Anotherpreferred material contains ZnO and Al₂O₃ in the form of a zincaluminate spinel. Yet another preferred material is a combination ofZnO, Al₂O₃, and CuO. Each of these materials may also contain anadditional metal function provided by a Group VIIIB metal, such as Pd orPt. Such metals may be preferred when a dehydrogenation reaction is tobe carried out in concert with the aldol condensation reaction. In oneembodiment, the base catalyst is a metal oxide containing Cu, Ni, Zn, V,Zr, or mixtures thereof. In another embodiment, the base catalyst is azinc aluminate metal containing Pt, Pd Cu, Ni, or mixtures thereof.

Preferred loading of the primary metal in the condensation catalyst isin the range of 0.10 wt % to 25 wt %, with weight percentages of 0.10%and 0.05% increments between, such as 1.00%, 1.10%, 1.15%, 2.00%, 2.50%,5.00%, 10.00%, 12.50%, 15.00% and 20.00%. The preferred atomic ratio ofthe second metal, if any, is in the range of 0.25-to-1 to 10-to-1,including ratios there between, such as 0.50, 1.00, 2.50, 5.00, and7.50-to-1.

In some embodiments, the base catalyzed condensation reaction isperformed using a condensation catalyst with both an acid and basefunctionality. The acid-aldol condensation catalyst may comprisehydrotalcite, zinc-aluminate, phosphate, Li, Na, K, Cs, B, Rb, Mg, Si,Ca, Sr, Ba, Al, Ce, La, Sc, Y, Zr, Ti, Zn, Cr, or any combinationthereof. In further embodiments, the acid-base catalyst may also includeone or more oxides from the group of Ti, Zr, V, Nb, Ta, Mo, Cr, W, Mn,Re, Al, Ga, In, Fe, Co, Ir, Ni, Si, Cu, Zn, Sn, Cd, P, and combinationsthereof. In an embodiment, the acid-base catalyst includes a metalfunctionality provided by Cu, Ag, Au, Pt, Ni, Fe, Co, Ru, Zn, Cd, Ga,In, Rh, Pd, Ir, Re, Mn, Cr, Mo, W, Sn, Os, alloys or combinationsthereof. In one embodiment, the catalyst further includes Zn, Cd orphosphate. In one embodiment, the condensation catalyst is a metal oxidecontaining Pd, Pt, Cu or Ni, and even more preferably an aluminate orzirconium metal oxide containing Mg and Cu, Pt, Pd or Ni. The acid-basecatalyst may also include a hydroxyapatite (HAP) combined with any oneor more of the above metals. The acid-base catalyst may beself-supporting or adhered to any one of the supports further describedbelow, including supports containing carbon, silica, alumina, zirconia,titania, vanadia, ceria, nitride, boron nitride, heteropolyacids, alloysand mixtures thereof.

In an embodiment, the condensation catalyst may also include zeolitesand other microporous supports that contain Group IA compounds, such asLi, NA, K, Cs and Rb. Preferably, the Group IA material is present in anamount less than that required to neutralize the acidic nature of thesupport. A metal function may also be provided by the addition of groupVIIIB metals, or Cu, Ga, In, Zn or Sn. In one embodiment, thecondensation catalyst is derived from the combination of MgO and Al₂O₃to form a hydrotalcite material. Another preferred material contains acombination of MgO and ZrO₂, or a combination of ZnO and Al₂O₃. Each ofthese materials may also contain an additional metal function providedby copper or a Group VIIIB metal, such as Ni, Pd, Pt, or combinations ofthe foregoing.

If a Group IIB, VIB, VIIB, VIIIB, IIA or IVA metal is included in thecondensation catalyst, the loading of the metal is in the range of 0.10wt % to 10 wt %, with weight percentages of 0.10% and 0.05% incrementsbetween, such as 1.00%, 1.10%, 1.15%, 2.00%, 2.50%, 5.00% and 7.50%,etc. If a second metal is included, the preferred atomic ratio of thesecond metal is in the range of 0.25-to-1 to 5-to-1, including ratiosthere between, such as 0.50, 1.00, 2.50 and 5.00-to-1.

The condensation catalyst may be self-supporting (i.e., the catalystdoes not need another material to serve as a support), or may require aseparate support suitable for suspending the catalyst in the reactantstream. One exemplary support is silica, especially silica having a highsurface area (greater than 100 square meters per gram), obtained bysol-gel synthesis, precipitation, or fuming. In other embodiments,particularly when the condensation catalyst is a powder, the catalystsystem may include a binder to assist in forming the catalyst into adesirable catalyst shape. Applicable forming processes includeextrusion, pelletization, oil dropping, or other known processes. Zincoxide, alumina, and a peptizing agent may also be mixed together andextruded to produce a formed material. After drying, this material iscalcined at a temperature appropriate for formation of the catalyticallyactive phase, which usually requires temperatures in excess of 452° C.Other catalyst supports as known to those of ordinary skill in the artmay also be used.

In some embodiments, a dehydration catalyst, a dehydrogenation catalyst,and the condensation catalyst can be present in the same reactor as thereaction conditions overlap to some degree. In these embodiments, adehydration reaction and/or a dehydrogenation reaction may occursubstantially simultaneously with the condensation reaction. In someembodiments, a catalyst may comprise active sites for a dehydrationreaction and/or a dehydrogenation reaction in addition to a condensationreaction. For example, a catalyst may comprise active metals for adehydration reaction and/or a dehydrogenation reaction along with acondensation reaction at separate sites on the catalyst or as alloys.Suitable active elements can comprise any of those listed above withrespect to the dehydration catalyst, dehydrogenation catalyst, and thecondensation catalyst. Alternately, a physical mixture of dehydration,dehydrogenation, and condensation catalysts could be employed. While notintending to be limited by theory, it is believed that using acondensation catalyst comprising a metal and/or an acid functionalitymay assist in pushing the equilibrium limited aldol condensationreaction towards completion. Advantageously, this can be used to effectmultiple condensation reactions with dehydration and/or dehydrogenationof intermediates, in order to form (via condensation, dehydration,and/or dehydrogenation) higher molecular weight oligomers as desired toproduce jet or diesel fuel.

The specific C4+ compounds produced in the condensation reaction willdepend on various factors, including, without limitation, the type ofoxygenated intermediates in the reactant stream, condensationtemperature, condensation pressure, the reactivity of the catalyst, andthe flow rate of the reactant stream as it affects the space velocity,GHSV and WHSV. Preferably, the reactant stream is contacted with thecondensation catalyst at a WHSV that is appropriate to produce thedesired hydrocarbon products. The WHSV is preferably at least about 0.1grams of oxygenated intermediates in the reactant stream per hour, morepreferably the WHSV is between about 0.1 to 40.0 g/g hr, including aWHSV of about 1, 2, 3, 4, 5, 6, 7, 8, 9, 10, 11, 12, 13, 14, 15, 20, 25,30, 35 g/g hr, and increments between.

In general, the condensation reaction should be carried out at atemperature at which the thermodynamics of the proposed reaction arefavorable. For condensed phase liquid reactions, the pressure within thereactor must be sufficient to maintain at least a portion of thereactants in the condensed liquid phase at the reactor inlet. For vaporphase reactions, the reaction should be carried out at a temperaturewhere the vapor pressure of the oxygenates is at least about 10 kPa, andthe thermodynamics of the reaction are favorable. The condensationtemperature will vary depending upon the specific oxygenatedintermediates used, but is generally in the range of from about 77° C.to 502° C. for reactions taking place in the vapor phase, and morepreferably from about 127° C. to 452° C. For liquid phase reactions, thecondensation temperature may be from about 7° C. to 477° C., and thecondensation pressure from about 0.1 kPa to 10,000 kPa. Preferably, thecondensation temperature is between about 17° C. and 302° C., or betweenabout 17° C. and 252° C. for difficult substrates.

Varying the factors above, as well as others, will generally result in amodification to the specific composition and yields of the C4+compounds. For example, varying the temperature and/or pressure of thereactor system, or the particular catalyst formulations, may result inthe production of C4+ alcohols and/or ketones instead of C4+hydrocarbons. The C4+ hydrocarbon product may also contain a variety ofolefins, and alkanes of various sizes (typically branched alkanes).Depending upon the condensation catalyst used, the hydrocarbon productmay also include aromatic and cyclic hydrocarbon compounds. The C4+hydrocarbon product may also contain undesirably high levels of olefins,which may lead to coking or deposits in combustion engines, or otherundesirable hydrocarbon products. In such event, the hydrocarbonmolecules produced may be optionally hydrogenated to reduce the ketonesto alcohols and hydrocarbons, while the alcohols and unsaturatedhydrocarbon may be reduced to alkanes, thereby forming a more desirablehydrocarbon product having low levels of olefins, aromatics or alcohols.

The condensation reactions may be carried out in any reactor of suitabledesign, including continuous-flow, batch, semi-batch or multi-systemreactors, without limitation as to design, size, geometry, flow rates,etc. The reactor system may also use a fluidized catalytic bed system, aswing bed system, fixed bed system, a moving bed system, or acombination of the above. In some embodiments, bi-phasic (e.g.,liquid-liquid) and tri-phasic (e.g., liquid-liquid-solid) reactors maybe used to carry out the condensation reactions.

In a continuous flow system, the reactor system can include an optionaldehydrogenation bed adapted to produce dehydrogenated oxygenatedintermediates, an optional dehydration bed adapted to produce dehydratedoxygenated intermediates, and a condensation bed to produce C4+compounds from the oxygenated intermediates. The dehydrogenation bed isconfigured to receive the reactant stream and produce the desiredoxygenated intermediates, which may have an increase in the amount ofcarbonyl-containing compounds. The de-hydration bed is configured toreceive the reactant stream and produce the desired oxygenatedintermediates. The condensation bed is configured to receive theoxygenated intermediates for contact with the condensation catalyst andproduction of the desired C4+ compounds. For systems with one or morefinishing steps, an additional reaction bed for conducting the finishingprocess or processes may be included after the condensation bed.

In an embodiment, the optional dehydration reaction, the optionaldehydrogenation reaction, the optional ketonization reaction, theoptional ring opening reaction, and the condensation reaction catalystbeds may be positioned within the same reactor vessel or in separatereactor vessels in fluid communication with each other. Each reactorvessel preferably includes an outlet adapted to remove the productstream from the reactor vessel. For systems with one or more finishingsteps, the finishing reaction bed or beds may be within the same reactorvessel along with the condensation bed or in a separate reactor vesselin fluid communication with the reactor vessel having the condensationbed.

In an embodiment, the reactor system also includes additional outlets toallow for the removal of portions of the reactant stream to furtheradvance or direct the reaction to the desired reaction products, and toallow for the collection and recycling of reaction byproducts for use inother portions of the system. In an embodiment, the reactor system alsoincludes additional inlets to allow for the introduction of supplementalmaterials to further advance or direct the reaction to the desiredreaction products, and to allow for the recycling of reaction byproductsfor use in other reactions.

In an embodiment, the reactor system also includes elements which allowfor the separation of the reactant stream into different componentswhich may find use in different reaction schemes or to simply promotethe desired reactions. For instance, a separator unit, such as a phaseseparator, extractor, purifier or distillation column, may be installedprior to the condensation step to remove water from the reactant streamfor purposes of advancing the condensation reaction to favor theproduction of higher hydrocarbons. In an embodiment, a separation unitis installed to remove specific intermediates to allow for theproduction of a desired product stream containing hydrocarbons within aparticular carbon number range, or for use as end products or in othersystems or processes.

The condensation reaction can produce a broad range of compounds withcarbon numbers ranging from C4 to C30 or greater. Exemplary compoundsinclude, but are not limited to, C4+ alkanes, C4+ alkenes, C5+cycloalkanes, C5+ cycloalkenes, aryls, fused aryls, C4+ alcohols, C4+ketones, and mixtures thereof. The C4+ alkanes and C4+ alkenes may rangefrom 4 to 30 carbon atoms (C4-C30 alkanes and C4-C30 alkenes) and may bebranched or straight chained alkanes or alkenes. The C4+ alkanes and C4+alkenes may also include fractions of C7-C14, C12-C24 alkanes andalkenes, respectively, with the C7-C14 fraction directed to jet fuelblend, and the C12-C24 fraction directed to a diesel fuel blend andother industrial applications. Examples of various C4+ alkanes and C4+alkenes include, without limitation, butane, butene, pentane, pentene,2-methylbutane, hexane, hexene, 2-methylpentane, 3-methylpentane,2,2-dimethylbutane, 2,3-dimethylbutane, heptane, heptene, octane,octene, 2,2,4,-trimethylpentane, 2,3-dimethyl hexane,2,3,4-trimethylpentane, 2,3-dimethylpentane, nonane, nonene, decane,decene, undecane, undecene, dodecane, dodecene, tridecane, tridecene,tetradecane, tetradecene, pentadecane, pentadecene, hexadecane,hexadecene, heptyldecane, heptyldecene, octyldecane, octyldecene,nonyldecane, nonyldecene, eicosane, eicosene, uneicosane, uneicosene,doeicosane, doeicosene, trieicosane, trieicosene, tetraeicosane,tetraeicosene, and isomers thereof.

The C5+ cycloalkanes and C5+ cycloalkenes have from 5 to 30 carbon atomsand may be unsubstituted, mono-substituted or multi-substituted. In thecase of mono-substituted and multi-substituted compounds, thesubstituted group may include a branched C3+ alkyl, a straight chain C1+alkyl, a branched C3+ alkylene, a straight chain C1+ alkylene, astraight chain C2+ alkylene, a phenyl or a combination thereof. In oneembodiment, at least one of the substituted groups include a branchedC3-C12 alkyl, a straight chain C1-C12 alkyl, a branched C3-C12 alkylene,a straight chain C1-C12 alkylene, a straight chain C2-C12 alkylene, aphenyl or a combination thereof. In yet another embodiment, at least oneof the substituted groups includes a branched C3-C4 alkyl, a straightchain C1-C4 alkyl, a branched C3-C4 alkylene, a straight chain C1-C4alkylene, a straight chain C2-C4 alkylene, a phenyl, or any combinationthereof. Examples of desirable C5+ cycloalkanes and C5+ cycloalkenesinclude, without limitation, cyclopentane, cyclopentene, cyclohexane,cyclohexene, methyl-cyclopentane, methyl-cyclopentene,ethyl-cyclopentane, ethyl-cyclopentene, ethyl-cyclohexane,ethyl-cyclohexene, and isomers thereof.

Aryls will generally consist of an aromatic hydrocarbon in either anunsubstituted (phenyl), mono-substituted or multi-substituted form. Inthe case of mono-substituted and multi-substituted compounds, thesubstituted group may include a branched C3+ alkyl, a straight chain C1+alkyl, a branched C3+ alkylene, a straight chain C2+ alkylene, a phenylor a combination thereof. In one embodiment, at least one of thesubstituted groups includes a branched C3-C12 alkyl, a straight chainC1-C12 alkyl, a branched C3-C12 alkylene, a straight chain C2-C12alkylene, a phenyl, or any combination thereof. In yet anotherembodiment, at least one of the substituted groups includes a branchedC3-C4 alkyl, a straight chain C1-C4 alkyl, a branched C3-C4 alkylene,straight chain C2-C4 alkylene, a phenyl, or any combination thereof.Examples of various aryls include, without limitation, benzene, toluene,xylene (dimethylbenzene), ethyl benzene, para xylene, meta xylene, orthoxylene, C9 aromatics.

Fused aryls will generally consist of bicyclic and polycyclic aromatichydrocarbons, in either an unsubstituted, mono-substituted ormulti-substituted form. In the case of mono-substituted andmulti-substituted compounds, the substituted group may include abranched C3+ alkyl, a straight chain C1+ alkyl, a branched C3+ alkylene,a straight chain C2+ alkylene, a phenyl or a combination thereof. Inanother embodiment, at least one of the substituted groups includes abranched C3-C4 alkyl, a straight chain C1-C4 alkyl, a branched C3-C4alkylene, a straight chain C2-C4 alkylene, a phenyl, or any combinationthereof. Examples of various fused aryls include, without limitation,naphthalene, anthracene, tetrahydronaphthalene, anddecahydronaphthalene, indane, indene, and isomers thereof.

The moderate fractions, such as C7-C14, may be separated for jet fuel,while heavier fractions, (e.g., C12-C24), may be separated for dieseluse. The heaviest fractions may be used as lubricants or cracked toproduce additional gasoline and/or diesel fractions. The C4+ compoundsmay also find use as industrial chemicals, whether as an intermediate oran end product. For example, the aryls toluene, xylene, ethyl benzene,para xylene, meta xylene, ortho xylene may find use as chemicalintermediates for the production of plastics and other products.Meanwhile, the C9 aromatics and fused aryls, such as naphthalene,anthracene, tetrahydronaphthalene, and decahydronaphthalene, may finduse as solvents in industrial processes.

In an embodiment, additional processes are used to treat the fuel blendto remove certain components or further conform the fuel blend to adiesel or jet fuel standard. Suitable techniques include hydrotreatingto reduce the amount of or remove any remaining oxygen, sulfur, ornitrogen in the fuel blend. The conditions for hydrotreating ahydrocarbon stream are known to one of ordinary skill in the art.

In an embodiment, hydrogenation is carried out in place of or after thehydrotreating process to saturate at least some olefinic bonds. In someembodiments, a hydrogenation reaction may be carried out in concert withthe aldol condensation reaction by including a metal functional groupwith the aldol condensation catalyst. Such hydrogenation may beperformed to conform the fuel blend to a specific fuel standard (e.g., adiesel fuel standard or a jet fuel standard). The hydrogenation of thefuel blend stream can be carried out according to known procedures,either with the continuous or batch method. The hydrogenation reactionmay be used to remove a remaining carbonyl group or hydroxyl group. Insuch event, any one of the hydrogenation catalysts described above maybe used. Such catalysts may include any one or more of the followingmetals, Cu, Ni, Fe, Co, Ru, Pd, Rh, Pt, Ir, Os, alloys or combinationsthereof, alone or with promoters such as Au, Ag, Cr, Zn, Mn, Sn, Cu, Bi,and alloys thereof, may be used in various loadings ranging from about0.01 wt % to about 20 wt % on a support as described above. In general,the finishing step is carried out at finishing temperatures of betweenabout 80° C. to 250° C., and finishing pressures in the range of about 5to 150 bar. In one embodiment, the finishing step is conducted in thevapor phase or liquid phase, and uses in situ generated H₂ (e.g.,generated in the HYDROGENOLYSIS reaction step), external H₂, recycledH₂, or combinations thereof, as necessary.

In an embodiment, isomerization is used to treat the fuel blend tointroduce a desired degree of branching or other shape selectivity to atleast some components in the fuel blend. It may be useful to remove anyimpurities before the hydrocarbons are contacted with the isomerizationcatalyst. The isomerization step comprises an optional stripping step,wherein the fuel blend from the oligomerization reaction may be purifiedby stripping with water vapor or a suitable gas such as lighthydrocarbon, nitrogen or hydrogen. The optional stripping step iscarried out in a counter-current manner in a unit upstream of theisomerization catalyst, wherein the gas and liquid are contacted witheach other, or before the actual isomerization reactor in a separatestripping unit utilizing counter-current principle.

After the optional stripping step the fuel blend can be passed to areactive isomerization unit comprising one or several catalyst bed(s).The catalyst beds of the isomerization step may operate either inco-current or counter-current manner. In the isomerization step, thepressure may vary from 2000 kPa to 15,000 kPa, preferably in the rangeof 2000 kPa to 10,000 kPa, the temperature being between 197° C. and502° C., preferably between 302° C. and 402° C. In the isomerizationstep, any isomerization catalysts known in the art may be used. Suitableisomerization catalysts can contain molecular sieve and/or a metal fromGroup VII and/or a carrier. In an embodiment, the isomerization catalystcontains SAPO-11 or SAPO41 or ZSM-22 or ZSM-23 or ferrierite and Pt, Pdor Ni and Al₂O₃ or SiO₂. Typical isomerization catalysts are, forexample, Pt/SAPO-11/Al₂O₃, Pt/ZSM-22/Al₂O₃, Pt/ZSM-23/Al₂O₃ andPt/SAPO-11/SiO₂.

Other factors, such as the concentration of water or undesiredoxygenated intermediates, may also effect the composition and yields ofthe C4+ compounds, as well as the activity and stability of thecondensation catalyst. In such event, the process may include adewatering step that removes a portion of the water prior to thecondensation reaction and/or the optional dehydration reaction, or aseparation unit for removal of the undesired oxygenated intermediates.For instance, a separator unit, such as a phase separator, extractor,purifier or distillation column, may be installed prior to thecondensation step so as to remove a portion of the water from thereactant stream containing the oxygenated intermediates. A separationunit may also be installed to remove specific oxygenated intermediatesto allow for the production of a desired product stream containinghydrocarbons within a particular carbon range, or for use as endproducts or in other systems or processes.

Thus, in one embodiment, the fuel blend produced by the processesdescribed herein is a hydrocarbon mixture that meets the requirementsfor jet fuel (e.g., conforms with ASTM D1655). In another embodiment,the product of the processes described herein is a hydrocarbon mixturethat comprises a fuel blend meeting the requirements for a diesel fuel(e.g., conforms with ASTM D975).

Yet in another embodiment of the invention, the C2+ olefins are producedby catalytically reacting the oxygenated intermediates in the presenceof a dehydration catalyst at a dehydration temperature and dehydrationpressure to produce a reaction stream comprising the C2+ olefins. TheC2+ olefins comprise straight or branched hydrocarbons containing one ormore carbon-carbon double bonds. In general, the C2+ olefins containfrom 2 to 8 carbon atoms, and more preferably from 3 to 5 carbon atoms.In one embodiment, the olefins comprise propylene, butylene, pentylene,isomers of the foregoing, and mixtures of any two or more of theforegoing. In another embodiment, the C2+ olefins include C4+ olefinsproduced by catalytically reacting a portion of the C2+ olefins over anolefin isomerization catalyst. In an embodiment, a method of forming afuel blend from a biomass feedstock may comprise a digester thatreceives a biomass feedstock and a digestive solvent operating underconditions to effectively remove nitrogen and sulfur compounds from saidbiomass feedstock and discharges a treated stream comprising acarbohydrate having less than 35% of the sulfur content and less than35% of the nitrogen content of the untreated biomass feedstock on a drymass basis; an hydrogenolysis and hydrodeoxygenation reactor comprisingan hydrogenolysis and hydrodeoxygenation catalyst that receives thetreated stream and discharges an oxygenated intermediate, wherein afirst portion of the oxygenated intermediate stream is recycled to thedigester as at least a portion of the digestive solvent; a first fuelsprocessing reactor comprising a dehydrogenation catalyst that receives asecond portion of the oxygenated intermediate stream and discharges anolefin-containing stream; and a second fuels processing reactorcomprising an alkylation catalyst that receives the olefin-containingstream and discharges a liquid fuel.

The dehydration catalyst comprises a member selected from the groupconsisting of an acidic alumina, aluminum phosphate, silica-aluminaphosphate, amorphous silica-alumina, aluminosilicate, zirconia, sulfatedzirconia, tungstated zirconia, tungsten carbide, molybdenum carbide,titania, sulfated carbon, phosphated carbon, phosphated silica,phosphated alumina, acidic resin, heteropolyacid, inorganic acid, and acombination of any two or more of the foregoing. In one embodiment, thedehydration catalyst further comprises a modifier selected from thegroup consisting of Ce, Y, Sc, La, Li, Na, K, Rb, Cs, Mg, Ca, Sr, Ba, P,B, Bi, and a combination of any two or more of the foregoing. In anotherembodiment, the dehydration catalyst further comprises an oxide of anelement, the element selected from the group consisting of Ti, Zr, V,Nb, Ta, Mo, Cr, W, Mn, Re, Al, Ga, In, Fe, Co, Ir, Ni, Si, Cu, Zn, Sn,Cd, P, and a combination of any two or more of the foregoing. In yetanother embodiment, the dehydration catalyst further comprises a metalselected from the group consisting of Cu, Ag, Au, Pt, Ni, Fe, Co, Ru,Zn, Cd, Ga, In, Rh, Pd, Ir, Re, Mn, Cr, Mo, W, Sn, Os, an alloy of anytwo or more of the foregoing, and a combination of any two or more ofthe foregoing.

In yet another embodiment, the dehydration catalyst comprises analuminosilicate zeolite. In one version, the dehydration catalystfurther comprises a modifier selected from the group consisting of Ga,In, Zn, Fe, Mo, Ag, Au, Ni, P, Sc, Y, Ta, a lanthanide, and acombination of any two or more of the foregoing. In another version, thedehydration catalyst further comprises a metal selected from the groupconsisting of Cu, Ag, Au, Pt, Ni, Fe, Co, Ru, Zn, Cd, Ga, In, Rh, Pd,Ir, Re, Mn, Cr, Mo, W, Sn, Os, an alloy of any two or more of theforegoing, and a combination of any two or more of the foregoing.

In another embodiment, the dehydration catalyst comprises a bifunctionalpentasil ring-containing aluminosilicate zeolite. In one version, thedehydration catalyst further comprises a modifier selected from thegroup consisting of Ga, In, Zn, Fe, Mo, Ag, Au, Ni, P, Sc, Y, Ta, alanthanide, and a combination of any two or more of the foregoing. Inanother version, the dehydration catalyst further comprises a metalselected from the group consisting of Cu, Ag, Au, Pt, Ni, Fe, Co, Ru,Zn, Cd, Ga, In, Rh, Pd, Ir, Re, Mn, Cr, Mo, W, Sn, Os, an alloy of anytwo or more of the foregoing, and a combination of any two or more ofthe foregoing.

The dehydration reaction is conducted at a temperature and pressurewhere the thermodynamics are favorable. In general, the reaction may beperformed in the vapor phase, liquid phase, or a combination of both. Inone embodiment, the dehydration temperature is in the range of about100° C. to 500° C., and the dehydration pressure is in the range ofabout 0 to 100 bar. In another embodiment, the dehydration temperatureis in the range of about 125° C. to 450° C., and the dehydrationpressure is at least 0.1 bar absolute. In another version, thedehydration temperature is in the range of about 150° C. to 350° C., andthe dehydration pressure is in the range of about 5 to 50 bar. In yetanother version, the dehydration temperature is in the range of about175° C. to 325° C.

The C6+ paraffins are produced by catalytically reacting the C2+ olefinswith a stream of C4+ isoparaffins in the presence of an alkylationcatalyst at an alkylation temperature and alkylation pressure to producea product stream comprising C6+ paraffins. The C4+ isoparaffins includealkanes and cycloalkanes having 4 to 7 carbon atoms, such as isobutane,isopentane, naphthenes, and higher homologues having a tertiary carbonatom (e.g., 2-methylbutane and 2,4-dimethylpentane), isomers of theforegoing, and mixtures of any two or more of the foregoing. In oneembodiment, the stream of C4+ isoparaffins comprises of internallygenerated C4+ isoparaffins, external C4+ isoparaffins, recycled C4+isoparaffins, or combinations of any two or more of the foregoing.

The C6+ paraffins will generally be branched paraffins, but may alsoinclude normal paraffins. In one version, the C6+ paraffins comprises amember selected from the group consisting of a branched C6-10 alkane, abranched C6 alkane, a branched C7 alkane, a branched C8 alkane, abranched C9 alkane, a branched C₁₀ alkane, or a mixture of any two ormore of the foregoing. In one version, the C6+ paraffins comprisedimethylbutane, 2,2-dimethylbutane, 2,3-dimethylbutane, methylpentane,2-methylpentane, 3-methylpentane, dimethylpentane, 2,3-dimethylpentane,2,4-dimethylpentane, methylhexane, 2,3-dimethylhexane,2,3,4-trimethylpentane, 2,2,4-trimethylpentane, 2,2,3-trimethylpentane,2,3,3-trimethylpentane, dimethylhexane, or mixtures of any two or moreof the foregoing.

The alkylation catalyst comprises a member selected from the group ofsulfuric acid, hydrofluoric acid, aluminum chloride, boron trifluoride,solid phosphoric acid, chlorided alumina, acidic alumina, aluminumphosphate, silica-alumina phosphate, amorphous silica-alumina,aluminosilicate, aluminosilicate zeolite, zirconia, sulfated zirconia,tungstated zirconia, tungsten carbide, molybdenum carbide, titania,sulfated carbon, phosphated carbon, phosphated silica, phosphatedalumina, acidic resin, heteropolyacid, inorganic acid, and a combinationof any two or more of the foregoing. The alkylation catalyst may alsoinclude a mixture of a mineral acid with a Friedel-Crafts metal halide,such as aluminum bromide, and other proton donors.

In one embodiment, the alkylation catalyst comprises an aluminosilicatezeolite. In one version, the alkylation catalyst further comprises amodifier selected from the group consisting of Ga, In, Zn, Fe, Mo, Ag,Au, Ni, P, Sc, Y, Ta, a lanthanide, and a combination of any two or moreof the foregoing. In another version, the alkylation catalyst furthercomprises a metal selected from the group consisting of Cu, Ag, Au, Pt,Ni, Fe, Co, Ru, Zn, Cd, Ga, In, Rh, Pd, Ir, Rr, Mn, Cr, Mo, W, Sn, Os,an alloy of any two or more of the foregoing, and a combination of anytwo or more of the foregoing.

In another embodiment, the alkylation catalyst comprises a bifunctionalpentasil ring-containing aluminosilicate zeolite. In one version, thealkylation catalyst further comprises a modifier selected from the groupconsisting of Ga, In, Zn, Fe, Mo, Ag, Au, Ni, P, Sc, Y, Ta, alanthanide, and a combination of any two or more of the foregoing. Inanother version, the alkylation catalyst further comprises a metalselected from the group consisting of Cu, Ag, Au, Pt, Ni, Fe, Co, Ru,Zn, Cd, Ga, In, Rh, Pd, Ir, Re, Mn, Cr, Mo, W, Sn, Os, an alloy of anytwo or more of the foregoing, and a combination of any two or more ofthe foregoing. In one version, the dehydration catalyst and thealkylation catalyst are atomically identical.

The alkylation reaction is conducted at a temperature where thethermodynamics are favorable. In general, the alkylation temperature isin the range of about −20° C. to 300° C., and the alkylation pressure isin the range of about 0 psig to 200 bar. In one version, the alkylationtemperature is in the range of about 100° C. to 300° C. In anotherversion, the alkylation temperature is in the range of about 0° C. to100° C., and the alkylation pressure is at least 5 bar. In yet anotherversion, the alkylation temperature is in the range of about 0° C. to50° C. and the alkylation pressure is less than 20 bar. In still yetanother version, the alkylation temperature is in the range of about 70°C. to 250° C., and the alkylation pressure is in the range of about 5 to100 bar. In one embodiment, the alkylation catalyst comprises a mineralacid or a strong acid and the alkylation temperature is less than ° C.In another embodiment, the alkylation catalyst comprises a zeolite andthe alkylation temperature is greater than 100° C.

Another aspect of the present invention is that the C4+ isoparaffins maybe generated internally by catalytically reacting an isoparaffinfeedstock stream comprising C4+ normal paraffins, aromatics and/ornaphthenes in the presence of an isomerization catalyst at anisomerization temperature and isomerization pressure to produceinternally generated C4+ isoparaffins. The C4+ normal paraffins willgenerally include alkanes having 4 to 7 carbon atoms, such as n-butane,n-pentane, n-hexane, n-heptane, and mixtures of any two or more of theforegoing. In one arrangement, the isoparaffin feedstock stream iscollected upstream of the alkylation catalyst from the reaction streamhaving the oxygenated intermediates or the reaction stream having theC2+ olefins and processed for the production of the internally generatedC4+ isoparaffins. In another arrangement, the C4+ normal paraffins,aromatics and/or naphthenes are collected downstream of the alkylationcatalyst from the product stream having the C6+ paraffins and thenrecycled for use in the production of the internally generated C₄₊isoparaffins. The C4+ isoparaffins may also be provided solely from anexternal source or used to supplement the internally generated C4+isoparaffins. In another version, the C4+ isoparaffins are recycled C4+isoparaffins collected from the product stream having the C6+ paraffins.

The isomerization catalyst is a catalyst capable of reacting a C4+normal paraffin, aromatic or naphthene to produce a C4+ isoparaffin. Inone version, the isomerization catalyst includes a zeolite, zirconia,sulfated zirconia, tungstated zirconia, alumina, silica-alumina, zincaluminate, chlorided alumina, phosphoric acid, or mixtures of any two ormore of the foregoing. In another version, the isomerization catalyst isan acidic beta, mordenite, or ZSM-5 zeolite. In yet another version, theisomerization catalyst further comprises a metal selected from the groupconsisting of Y, Pt, Ru, Ad, Ni, Rh, Ir, Fe, Co, Os, Zn, a lanthanide,or an alloy or combination of any two or more of the foregoing. In stillyet another version, the isomerization catalyst comprises a support, thesupport comprising alumina, sulfated oxide, clay, silica gel, aluminumphosphate, bentonite, kaolin, magnesium silicate, magnesium carbonate,magnesium oxide, aluminum oxide, activated alumina, bauxite, silica,silica-alumina, activated carbon, pumice, zirconia, titania, zirconium,titanium, kieselguhr, or zeolites.

In an embodiment of the present invention, the fuel yield of the currentprocess may be greater than other bio-based feedstock conversionprocesses. Without wishing to be limited by theory, it is believed thatsubstantially removing nitrogen compounds and sulfur compounds from thesoluble carbohydrate prior to the direct hydrogenolysis allows for agreater percentage of the biomass to be converted into higherhydrocarbons while limiting the formation of degradation products.

To facilitate a better understanding of the present invention, thefollowing examples of certain aspects of some embodiments are given. Inno way should the following examples be read to limit, or define, theentire scope of the invention.

EXAMPLES

Reaction studies were conducted in a Parr5000 Hastelloy multireactorcomprising 6×75-milliliter reactors operated in parallel at pressures upto 135 bar, and temperatures up to 275° C., stirred by magnetic stirbar. Alternate studies were conducted in 100-ml Parr4750 reactors, withmixing by top-driven stir shaft impeller, also capable of 135 bar and275° C. Larger scale extraction, pretreatment and digestion tests wereconducted in a 1-Liter Parr reactor with annular basket housing biomassfeed, or with filtered dip tube for direct contacting of biomassslurries.

Reaction samples were analyzed for sugar, polyol, and organic acidsusing an HPLC method entailing a Bio-Rad Aminex HPX-87H column (300mm×7.8 mm) operated at 0.6 ml/minute of a mobile phase of 5 mM SulfuricAcid in water, at an oven temperature of 30° C., a run time of 70minutes, and both RI and UV (320 nm) detectors.

Product formation (mono-oxygenates, glycols, diols, alkanes, acids) weremonitored via a gas chromatographic (GC) method “DB5-ox”, entailinga60-m×0.32 mm ID DB-5 column of 1 um thickness, with 50:1 split ratio, 2ml/min helium flow, and column oven at 40° C. for 8 minutes, followed byramp to 285° C. at 10° C./min, and a hold time of 53.5 minutes. Injectortemperature is set at 250° C., and detector temperature at 300° C.

Gasoline production potential by condensation was assessed via injectionof one microliters of liquid intermediate product into a catalytic pulsemicroreactor entailing a GC insert packed with 0.12 grams of ZSM-5catalyst, held at 375° C., followed by Restek Rtx-1701 (60-m) and DB-5(60-m) capillary GC columns in series (120-m total length, 0.32 mm ID,0.25 um film thickness) for an Agilent/HP 6890 GC equipped with flameionization detector. Helium flow was 2.0 ml/min (constant flow mode),with a 10:1 split ratio. Oven temperature was held at 35° C. for 10minutes, followed by a ramp to 270° C. at 3° C./min, followed by a 1.67minute hold time. Detector temperature was 300° C.

Example 1 Nickel Catalyst

A 100-ml Parr reactor was charged with 0.9243-grams of Ni-5249P catalyst(64% nickel on silica; Strem Chemicals, Inc.), 3.59-grams of milledbagasse (1-mm grate, 5% moisture), and 60.1-grams of deionized water.The reactor was pressured to 510 psig (3600 kPa) with H₂, before heatingto 170 C for 2.5 hours, then 190 C for 2.5 hours, followed by heating to210° C. for 22 hours.

Two additional cycles were conducted with the same temperature sequence,following addition of 3.62 and 3.58 grams of bagasse. A final pH of 3.89was observed, indicating acid coproduct formation. Formation of 2.1 wt %acetic acid was verified via DB5-ox GC analysis. Following reaction,solids were recovered by filtration on Whatman #2 filter paper, and ovendried overnight at 90° C. to assess the extent of digestion of biomass.Greater than 98% of the total bagasse charged over three cycles wasdigested. Ethylene glycol (EG) and 1,2-propylene glycol (PG) comprisedmore than 26.4% of the hydrocarbon products, as measured via DB5-ox GCmethod (Table 1). The remainder of product analyzed as a mixture ofprimarily C₂-C₆ oxygenates (alcohols, ketones) and carboxylic acids,suitable for condensation to liquid biofuels.

TABLE 1 Bagasse Hydrogenolysis with Nickel catalyst wt % of HC Componentproducts Ethylene glycol 8.6 1,2-Propylene glycol 17.8 Glycerol 5.7Erythritol 4.2 Total polyols 36.4 Total glycols 26.4 Weight percent oftotal hydrocarbon (HC) products.

One microliter of liquid product was injected onto the ZSM-5 pulsemicroreactor at 370° C. at a 1:15 split ratio to assess gasolineformation potential. Formation of alkanes, benzene, toluene, xylenes,trimethlybenzenes, and naphthalenes were observed at an approximateyield of 59% relative to that expected from complete conversion of thecarbohydrate fraction of the feed bagasse. This result demonstratesco-production of glycols and liquid biofuels via direct hydrogenolysisof biomass, followed by acid-catalyzed condensation of oxygenatespresent in the hydrogenolysis product stream.

Example 2 Sulfided Cobalt Molybdate Catalyst

A multi-cycle experiment was conducted using a nominal 3.50 grams ofbagasse with 1.04 grams of sulfided cobalt-molybdate catalyst (DC-2533catalyst from Criterion Catalyst & Technologies L.P) and 58.50 grams ofdeionized water. The catalyst was sulfided by the method described inUS2010/0236988, Example 5. The Parr 100-ml reactor was pressured to 1024psig with H₂ (7200 kPa), and heated to 170° C., and ramped to 240° C.over 7 hours, before holding at 240° C. overnight to completed aninitial cycle. Four additional cycles were completed in subsequent24-hour periods, entailing 9-hour ramps from 160-250° C., before holdingat 250° C. overnight. A total of 17.59 grams of bagasse were charged forthe five cycles.

A final pH of 3.49 was measured, indicated acid formation from thebiomass feed. DB5-ox GC analysis indicated 1.67% acetic acid present inthe final reaction liquid. Following reaction, solids were recovered byfiltration on Whatman #2 filter paper, and oven dried overnight at 90°C. to assess the extent of digestion of biomoass. Results indicaed 73%of the total bagasse charged over was digested into liquid solubleproducts. Ethylene glycol (10.8%) and 1,2-propylene glycol (14.9%)comprised more than 25.7% of the hydrocarbon products, as measured viaDB5-ox GC method (Table 2). The remainder of product analyzed as amixture of primarily C2-C6 oxygenates (alcohols, ketones), andcarboxylic acids, suitable for condensation to liquid biofuels.

Liquid product was injected onto the ZSM-5 pulse microreactor at 375° C.to assess gasoline formation potential. Formation of alkanes, benzene,toluene, xylenes, trimethlybenzenes, and naphthalenes were observed atan approximate yield of 36% relative to that expected from completeconversion of the carbohydrate fraction of the feed bagasse. This resultdemonstrates co-production of glycols and liquid biofuels via directhydrogenolysis of biomass over sulfided cobalt-molybdate catalyst,followed by acid-catalyzed condensation of oxygenates present in thehydrogenolysis product stream.

TABLE 2 Bagasse Hydrogenolysis with Sulfided Cobalt-Molybdate catalystWt % of Total Component HC products Ethylene glycol 10.8 1,2-Propyleneglycol 14.9 Glycerol 6.6 Erythritol 11.7 Total polyols 44.0 Totalglycols 25.7

Example 3 Use of Calcium Carbonate Cocatalyst/Buffer

Example 2 was repeated with addition of 2.06 grams of calcium carbonatefor the initial reaction, followed by addition of 0.50-0.51 grams ofcalcium carbonate for each successive cycle, to maintain a pH of greaterthan 4.5 throughout the reaction sequence. A final pH of 4.84 wasmeasured at the end of the fifth cycle. A total of 18.71 grams ofbagasse (dry basis) were charged across the five reaction cycles.

Following reaction, solids were recovered by filtration on Whatman #2filter paper, and oven dried overnight at 90° C. to assess the extent ofdigestion of biomoass. Results indicated 90% of the total bagassecharged over was digested into liquid soluble products. Ethylene glycol(9.1%) and 1,2-propylene glycol (32.8%) comprised more than 41% of thehydrocarbon products, as measured via DB5-ox GC method (Table 3). Theremainder of product analyzed as a mixture of primarily C2-C6 oxygenates(alcohols, ketones), and carboxylic acids, suitable for condensation toliquid biofuels.

Liquid product was injected onto the ZSM-5 pulse microreactor at 375° C.to assess gasoline formation potential. Formation of alkanes, benzene,toluene, xylenes, trimethlybenzenes, and naphthalenes were observed atan approximate yield of 50% relative to that expected from completeconversion of the carbohydrate fraction of the feed bagasse. This resultdemonstrates co-production of glycols and liquid biofuels via directhydrogenolysis of biomass over sulfided cobalt-molybdate catalyst,followed by acid-catalyzed condensation of oxygenates present in thehydrogenolysis product stream. Use of a basic buffer such as calciumcarbonate to improve yields of glycols, and moderate pH, is alsoestablished.

TABLE 3 Hydrogenolysis with sulfided cobalt molybdate catalyst andcalcium carbonate buffer wt % of total Component HC products Ethyleneglycol 9.1 1,2-Propylene glycol 32.8 Glycerol 1.0 Erythritol 0.2 Totalpolyols 43.0 Total glycols 41.9

Example 4 Sulfided Cobalt Molybdate Catalyst with KOH Buffer

Experiment 2 was repeated with addition of 1N KOH to buffer pH to 5.5for each reaction step. Three reaction cycles were conducted withaddition of 10.03 grams of bagasse (dry basis). A final pH of 5.34 wasmeasured for the liquid product of three cycles.

Following reaction, solids were recovered by filtration on Whatman #2filter paper, and oven dried overnight at 90° C. to assess the extent ofdigestion of biomoass. Results indicated 87.9% of the total bagassecharged over was digested into liquid soluble products. Ethylene glycol(5.1%) and 1,2-propylene glycol (16.7%) comprised more than 21% of thehydrocarbon products, as measured via DB5-ox GC method (Table 4).Further conversion of glycerol (8.2%) to propylene glycol can beachieved via continuing the —OH hydrogenolysis reaction, resulting inhigher yields of glycol products. The remainder of product analyzed as amixture of primarily C2-C6 oxygenates (alcohols, ketones) and carboxylicacids, suitable for condensation to liquid biofuels.

Liquid product was injected onto the ZSM-5 pulse microreactor at 375° C.to assess gasoline formation potential. Formation of alkanes, benzene,toluene, xylenes, trimethlybenzenes, and naphthalenes were observed atan approximate yield of 69% relative to that expected from completeconversion of the carbohydrate fraction of the feed bagasse. This resultdemonstrates co-production of glycols and liquid biofuels via directhydrogenolysis of biomass over sulfided cobalt-molybdate catalyst,followed by acid-catalyzed condensation of oxygenates present in thehydrogenolysis product stream. Use of potassium hydroxide as a basicbuffer to maintain pH>5 was demonstrated to give high yields of glycolintermediate products.

TABLE 4 Bagasse Hydrogenolysis with Sulfided Cobalt Molybdate catalystand KOH buffer wt % of HC Component products Ethylene glycol 5.11,2-Propylene glycol 16.7 Glycerol 8.2 Erythritol 12.0 Total polyols42.0 Total glycols 21.8

Example 5 Hardwood Hydrogenolysis with Ruthenium Catalyst

2.037 grams of ground Northern hardwood were charged with 20.073 gramsof deionized water, and 0.470 grams of 5% Ru/C Escat 4401 catalyst(Strem Chemicals, 50% wet), to a 75-ml Parr5000 reactor. The reactor waspressured to 700 psig with H₂, and ramped from 170-240° C. over 6 hours,before maintaining 240° C. overnight to complete reaction. The sequencewas repeated for two additional cycles with addition of 1.99 and 2.09grams of ground hardwood. Final pH measured at the end of the reactioncycles was 3.20.

Following reaction, solids were recovered by filtration on Whatman #2filter paper, and oven dried overnight at 90° C. to assess the extent ofdigestion of biomoass. Results indicated 88% digestion of the hardwoodcharged, into liquid soluble products. Ethylene glycol (4.9%) and1,2-propylene glycol (18.1%) comprised more than 22% of the hydrocarbonproducts, as measured via DB5-ox GC method (Table 5).

TABLE 5 Hardwood hydrogenolysis with Ruthenium catalyst wt % of HCComponent products Ethylene glycol 4.9 1,2-Propylene glycol 18.1Glycerol 4.1 Erythritol 1.1 Total polyols 28.1 Total glycols 22.9

Example 6 Softwood (Pine) Hydrogenolysis with Ruthenium Catalyst

The experiment of example 5 was repeated with addition of groundSouthern pine as the biomass feed, and with use of a Re-promoted 1.9%Pt/zirconia catalyst at Re:Pt rato of 3.75:1 prepared according to themethod in Example 7 of US2008/0215391. 6.12 grams of pine wood wereadded over three cycles. Final pH measured after three cycles was 3.06.

Following reaction, solids were recovered by filtration on Whatman #2filter paper, and oven dried overnight at 90° C. to assess the extent ofdigestion of biomoass. Results indicated greater than 98% digestion ofthe hardwood charged, into liquid soluble products. Ethylene glycol(9.8%) and 1,2-propylene glycol (14.1%) comprised more than 22% of thehydrocarbon products, as measured via DB5-ox GC method (Table 6).Further conversion of glycerol (7.4%) to propylene glycol can beachieved via continuing the —OH hydrogenolysis reaction, to furtherincrease glycol yields. The remainder of product analyzed as a mixtureof primarily C2-C6 oxygenates (alcohols, ketones), and carboxylic acids,suitable for condensation to liquid biofuels.

Liquid product was injected onto the ZSM-5 pulse microreactor at 375° C.to assess gasoline formation potential. Formation of alkanes, benzene,toluene, xylenes, trimethlybenzenes, and naphthalenes were observed atan approximate yield of 30% relative to that expected from completeconversion of the carbohydrate fraction of the feed softwood.

TABLE 6 Softwood (Pine) Hydrogenolysis with Pt—Re/zirconia catalyst wt %of HC Component products Ethylene glycol 9.8 1,2-Propylene glycol 14.1Glycerol 7.4 Erythritol 1.1 Total polyols 32.4 Total glycols 23.9

Examples 7-25 Hydrogenolysis of Sorbitol Over Nickel, Ruthenium andCobalt Catalysts

Nickel(Strem Chemicals, Inc; Sigma-Aldrich Inc; x-Activated Metals Inc),ruthenium (Strem Chemicals, Inc), and cobalt (W.R. Grace & Co., RaneyType 2724) hydrogenolysis catalysts were examined for the hydrogenolysisof sorbitol, representative of the sugar alcohol formed from hydrolysisand hydrogenation of sugar, starch, or biomass feed to an aqueous phaserefoming or hydrogenolysis reactor. Conversion of sorbitol and formationof glycol and polyol products were determined by HPLC analysis. Resultsshow formation of glycol intermediates, with polyol yield (includingglycerol) of 30-70% over the range of conditions tested. Sodiumcarbonate and calcium hydroxide were used to buffer pH over the range of3-9.

FIG. 4 shows a cross plot of conversion of sorbitol, vs. selectivity toglycols including glycerol. Selectivity is improved at lower conversionsof sorbitol. Selectivities observed were never greater than about 70%,which demonstrates the value of the current invention where themonoxygenates-rich stream formed as the remainder of the reactionproducts, is converted to value added liquid fuels. Controllingconversion from about 50% to about 80% by variation in contact time withcatalyst, or temperature, allows variation in the amount of glycolsformed, relative to the monooxyenate intermediates which are convertedto liquid fuels. Production ratios of glycols to liquid fuels can thusbe controlled, to optimize commercial economics.

TABLE 7 Hydrogenolysis with 50% Sorbitol feed over various catalystsTotal Psi Final Time GLY EG PG Polyol Sorb Polyol Ex# Catalyst Base T CH2 pH hours wt % wt % wt % wt % conv Yield 7 Ru/C Na₂CO₃ 200 1175 8.8819.10 0.00 1.31 5.16 18.50 0.76 0.37 8 Ni5249P Na₂CO₃ 195 1000 8.8123.10 0.00 2.93 14.13 17.92 0.98 0.36 9 A7000 Na₂CO₃ 210 1010 5.84 22.001.11 1.72 4.04 15.27 0.83 0.31 10 Ru/C Na₂CO₃ 210 1015 7.99 23.10 0.001.52 7.09 11.14 0.95 0.22 11 RaCo2724 Na2CO₃ 210 1012 6.03 22.00 0.701.31 3.22 17.04 0.76 0.34 12 Ni5249P Na2CO₃ 210 1010 7.56 6.75 0.61 3.659.52 26.99 0.74 0.54 13 RaCo2724 Na2CO₃ 210 1005 5.12 23.10 1.51 1.213.43 27.95 0.56 0.56 14 RaNi Ca(OH)₂ 210 1300 4.11 23.80 1.52 2.13 4.8624.73 0.68 0.49 15 Ru/C Ca(OH)₂ 210 1009 4.90 19.60 2.71 3.70 11.5435.04 0.66 0.70 16 Ni5249P Na₂CO₃ 220 1272 5.31 23.60 0.58 3.84 17.2922.50 0.98 0.45 17 Ni5249P Ca(OH)₂ 210 1010 4.83 22.80 3.24 3.49 15.9524.82 0.96 0.50 18 Ni5249P Ca(OH)₂ 210 1266 3.96 7.70 8.03 3.89 10.3633.37 0.78 0.67 19 5% Ru/C Ca(OH)₂ 210 1270 3.68 22.50 4.12 1.60 3.6630.40 0.58 0.61 20 RaNi* Ca(OH)₂ 220 798 3.27 18.50 2.94 1.82 4.86 32.850.54 0.66 21 RaCo2724 Ca(OH)₂ 210 990 3.79 25.00 2.75 1.50 5.25 29.130.61 0.58 22 Ni5249P Ca(OH)₂ 210 1300 3.79 20.10 2.61 3.09 17.10 23.041.00 0.46 23 5% Ru/C Ca(OH)₂ 220 992 3.95 19.00 3.01 1.81 5.12 25.590.69 0.51 24 5% Ru/C Ca(OH)₂ 210 1309 3.71 19.10 4.47 1.83 5.08 28.260.66 0.57 25 Ni5249P Ca(OH)₂ 220 1255 3.82 7.40 4.45 3.24 14.98 25.890.94 0.52 Ru/C = 5% Ru/C Escat ™ 4401 catalyst (Strem Chemicals, Inc.50% wet) Ni5249P = 64% nickel on silica (Strem Chemicals, Inc.) A7000 =Mo-promoted sponge metal catalyst (Activated Metals Inc.) RaCo2724 = W RGrace & Co.: Raney cobalt RaNi = Mo-promoted Raney-type nickel catalyst(Aldrich Chemical Company Inc.) GLY = glycerol Sorb = sorbitol

Example 26 Separation of Glycols from Hydrogenolysis ReactionIntermediates

A short-path distillation was conducted with 71.184 grams of a highconversion flow reaction product obtained from the hydrogenolysis at 240C and 667 psig under co-feed of H₂, of 50% sorbitol overrhenium-promoted 1.9% Pt/zirconia catalyst. The flash distillation wasconducted at atmospheric pressure under nitrogen, with a 2:1 refluxratio. An initial distillation cut comprising 17.6% of the total feed,was distilled at a tops temperature of 77-79° C. An organic layer of13.3% by volume was observed in the distilled sample. Distillation wascontinued until no further offtake at a tops temperature of 101° C., anda bottoms temperature of 120° C. HPLC analysis of feed and final bottomssamples are shown in Table 8. Heavy glycols and unconverted sorbitolwere concentrated in the distillation bottoms, with low losses tooverheads. DB5-ox analysis of mono-oxygenates and glycols indicatedgreater than 88% rejection of monooxygenates into the overheadeddistillate relative to feed, with less than 10% of monooxygenate lightends of boiling point less than glycerol, present in final bottomsproduct. Further separation via via an increased number of stages can beexpected to give improved separation, if desired.

TABLE 8 Rash distillation of hydrogenolysis reactor product Sorbitol EGPG Gly Total Sample wt % wt % wt % wt % grams feed 5.37 1.84 1.39 1.0771.184 bottoms 18.57 2.93 5.87 3.95 19.960 conc ratio 3.46 1.59 4.233.6A9 3.566 Gly = glycerol

Example 27 Separation and Purification of Glycol Co-Products

An Aspen Plus (version 2006.5) simulation (FIG. 1) was conducted for amodel hydrogenolysis reaction outlet stream comprising 2-propanol,ethanol, ethylene glycol, propylene glycol, and unconverted sorbitol,simulating the crude product from hydrogenolysis of a bio-based feedstream. Separation of light alcohols, water, and glycols fromunconverted sorbitol as an overhead stream was readily effected by flashseparation at ambient pressure (1-5 psig). Unconverted sorbitol can berecycle to reaction. A partial flash separation with recycle of somewater, sorbitol, and glycols back to reaction is preferred, to preventsorbitol precipitation. Overheads from the flash separation comprisingethanol, 2-propanol, water, ethylene glycol, and propylene were routedto an atmospheric pressure distillation column (nominal 1-5 psig) toseparate ethanol, 2-propanol, and water from ethylene glycol andpropylene glycol. A 20 stage column operated at a reflux ratio of 4,with 100° C. top distillate temperature, and 189° C. bottomstemperature, effected virtual complete separation of water and lightmonooxygenates, from the ethylene glycol (EG) and propylene glycol (PG)product mixture. The bottoms glycol mixture from this column comprisingethylene glycol and propylene glycol could be fractionated to give a125° C. tops distillate stream comprising 1,2-propylene glycol atgreater than 99% purity, and a 139° C. bottoms stream comprisingethylene glycol at greater than 99% purity, via a 90 stage columnoperating at a reflux ratio of 10, and under moderate vacuum of 100 tonabsolute pressure.

While many different separation sequences can be developed by thoseskilled in the art, the current example shows one configuration wherepurified ethylene glycol and propylene glycol product streams may beobtained, as separate products.

What is claimed is:
 1. A process of co-producing biofuels and glycolscomprising: (i) providing a bio-based feedstock stream containingcarbohydrates and water; (ii) contacting, in a first reaction system,the bio-based feedstock stream with hydrogen in the presence of ahydrogenolysis catalyst at a temperature in the range of 120° C. to 280°C. and 0.1 to 150 bar of hydrogen to produce a hydrogenolysis streamcontaining at least 5 wt %, based on the total oxygenates content, ofglycols that comprises ethylene glycol (EG) and 1,2-propylene glycol(PG), and other monooxygenates; (iii) contacting, in a second reactionsystem, at least a first portion of said hydrogenolysis stream withhydrogen in the presence of a hydrogenolysis catalyst at a temperaturein the range of 160° C. to 280° C. and in the presence of 0.1 to 150barhydrogen to produce an oxygenated intermediate stream comprising atleast 5 wt %, based on the total oxygenates, of monooxygenatedhydrocarbons of chain length less than 6 carbons; (iv) processing atleast a portion of the oxygenated intermediate stream to form a liquidfuel; (v) providing a second portion of said hydrogenolysis stream to afirst separation system; (vi) separating a portion of saidhydrogenolysis stream, in the first separation system, to amonooxygenates stream comprising monooxygenates and a glycol rich streamcomprising at least 10 wt %, based on the total oxygenates, of glycolsby flashing; and (vii) recovering glycols from the glycol rich stream.2. The process of claim 1 wherein further comprising (viii) recycling atleast a portion of said monooxygenates stream to the first reactionsystem.
 3. The process of claim 1 wherein glycols are recovered byseparating finished glycol from the glycol rich stream.
 4. The processof claim 1 wherein the glycol content of the hydrogenolysis streamproduced in step (ii) is at least 15 wt % based on the total oxygenatescontent.
 5. The method process of claim 1 wherein the ratio of thehydrogenolysis stream provided to the first separation system and to thesecond reaction system is in the range of 1.5:1 to 10:1.
 6. The processof claim 1 wherein the maximum steady state temperature of the firstreaction system is at least 10° C. less than the maximum steady statetemperature of the second reaction system;
 7. The process of claim 6wherein no process heat from external source is added between the firstand second reaction systems, or to the second reaction system, whereinthe temperature increase in the second reaction system results from theexothermic heats from the first reaction system and the second reactionsystem.
 8. The process of claim 4 wherein the glycol rich stream in step(vi) comprises at least 25 wt %, based on the total oxygenates, ofglycols.
 9. The process of claim 1 wherein bio-based feedstock issubjected to a digestive solvent to provide bio-based feedstock streamcomprising carbohydrates and water.
 10. The process of claim 1 whereinthe conversion of carbohydrate in the first reaction system is limitedto less than 80%.
 11. The method process of claim 10 wherein theconversion of carbohydrate in the first reaction system is limited toless than 70%.
 12. The process of claim 11 wherein the conversion ofcarbohydrate in the first reaction system is limited to less than 60%.13. The method process of claim 12 wherein the conversion ofcarbohydrate in the first reaction system is limited to less than 50%.14. The process of claim 1 wherein the conversion of carbohydrates tomonooxygenated compounds across the first reaction system and the secondreaction system is less than 90%.
 15. The process of claim 14 whereinthe oxygenated intermediate stream from the second reaction system isflash distilled to provide a stream containing at least a portion ofmonooxygenate intermediates as feed for production of liquid fuels. 16.The process of claim 15 wherein the bottoms from flashing containingenriched concentrations of unconverted carbohydrates, sugar alcohols,glycols, and some monooxygenates, are recycle to the first reactionsystem.
 17. A method process of co-producing biofuels and glycolscomprising: (i) providing a bio-based feedstock; (ii) contacting thebio-based feedstock with a digestive solvent to provide a bio-basedfeedstock stream containing carbohydrates and water; (iii) contacting,in a first reaction system, the bio-based feedstock stream with hydrogenin the presence of a hydrogenolysis catalyst at a temperature in therange of 120° C. to 280° C. and 0.1 to 150 bar of hydrogen to produce ahydrogenolysis stream containing at least 5 wt % , based on the totaloxygenates content, of glycols that comprises ethylene glycol (EG) and1,2-propylene glycol (PG), and other monooxygenates; (iv) contacting, ina second reaction system, at least a first portion of saidhydrogenolysis stream with hydrogen in the presence of a hydrogenolysiscatalyst at a temperature in the range of 160 ° C. to 280 ° C. and inthe presence of 0.1 to 150bar hydrogen to produce an oxygenatedintermediate stream comprising at least 5 wt %, based on the totaloxygenates, of monooxygenated hydrocarbons of chain length less than 6carbons; (v) processing at least a portion of the oxygenatedintermediate stream to form a liquid fuel; (vi) providing a secondportion of said hydrogenolysis stream to a first separation system;(vii) separating a portion of said hydrogenolysis stream, in the firstseparation system, to a monooxygenates stream comprising monooxygenatesand a glycol rich stream comprising at least 10 wt %, based on the totaloxygenates, of glycols by flashing; and (viii) recovering glycols fromthe glycol rich stream.
 18. The process of claim 17 wherein furthercomprising (ix) recycling at least a portion of said monooxygenatesstream to the first reaction system.
 19. The process of claim 17 whereinglycols are recovered by separating finished glycol from the glycol richstream.
 20. The process of claim 17 wherein the oxygenated intermediatesis subjected to condensation to produce a liquid fuel.
 21. The processof claim 17 wherein the oxygenated intermediates is subjected todehydration and alkylation to produce a liquid fuel.
 22. A process ofco-producing biofuels and glycols comprising: (i) providing a bio-basedfeedstock stream containing carbohydrates and water; (ii) contacting, ina first reaction zone, the bio-based feedstock stream with hydrogen inthe presence of a hydrogenolysis catalyst at a temperature in the rangeof 120° C. to 280° C. and 0.1 to 150 bar of hydrogen to produce ahydrogenolysis intermediate containing at least 5 wt %, based on thetotal oxygenates content, of glycols that comprises ethylene glycol (EG)and 1,2-propylene glycol (PG), and other monooxygenates; (iii)contacting, in a second reaction zone, said hydrogenolysis intermediatewith hydrogen in the presence of a hydrogenolysis catalyst at atemperature in the range of 160° C. to 280° C. and in the presence of0.1 to 150 bar hydrogen to produce a combined glycol and oxygenatedintermediate stream comprising at least 5 wt %, based on the totaloxygenates, of monooxygenated hydrocarbons of chain length less than 6carbons, and greater than 5 wt % glycols; (iv) separating, by flashing,the combined glycol and oxygenated intermediate stream into a glycolrich stream comprising at least 10 wt %, based on the total oxygenates,of glycols, and a mono-oxygenates-rich stream; (v) processing at least aportion of the monooxygenates-rich stream to form a liquid fuel; and(vi) recovering glycols from the glycol rich stream.
 23. The process ofclaim 22 where the temperature of the second reaction zone is at least10° C. higher than the temperature of the first reaction zone.
 24. Theprocess of claim 22 where at portion of either or both of theglycol-rich stream or mono-oxygenates rich stream is recycled to providesolvent for the first reaction zone.
 25. The process of claim 24 wherethe recycled solvent is routed to a digestion zone to digest solidbiomass to provide feed for the first reaction zone.